PROCESS DESIGN MANUAL
FLARE SYSTEM
670-225-9048
PLEASE READ MESSAGE BELOW BEFORE YOU
PROCEED
It is recognized that this manual will require further improvement and updating. However, this manual
may be used in actual projects with suitable caution. A listing of recommended action items is
provided on the next page to identify issues, which have been targeted as areas for improvement. A
brief review of those items is warranted before using the manual. Any suggestions for improvement
are welcomed and should be forwarded to David Kang.
The risk based assessment and recommendations of relief load mitigation instrumentation in Sections
5.3.4 and 5.3.5 shall be used only as a reference. The values of reliability and/or probability used in
these sections require further verification and approval based on the latest industry practices. The
recommendations for use of load mitigation instruments may be changed depending on the finalized
instrument reliability values.
Please E-Mail David Kang of Process Engineering in the Irvine office for questions or comments
regarding this Design Manual.
ISSUE
0
1
Cover Page.doc
DATE
11-Aug-98
14-Apr-00
MADE
CAS
YVB
CHECKED
DKK
DKK
APPROVED
DDC
DDC
DESCRIPTION
Review & Comment
Issuance
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FLARE SYSTEM
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AREAS FOR FURTHER IMPROVEMENTS
SECTION 3.11.5.4 TWO PHASE OR FLASHING FLOW SERVICE
Provide the methodology for using Diers HEM to calculate relief valve size for two phase or
flashing flow relief. This method has been recently adopted by API.
SECTION 5.3.4 THRU 5.3.5 RISK BASED ASSESSMENT AND RECOMMENDATIONS FOR
RELIEF LOAD MITIGATION INSTRUMENTATION
Update reliability and probability values for relief load mitigation instruments, and establish
standards for use of the load mitigation instruments and mitigated flare load calculation
methodology. The Fluor Daniel Houston office is currently (1998) reviewing Fluor Daniel’s
position on this subject.
SECTION 7.5.1.3 WATER SEALS
Review number of contraints on seal drum criteria depicted in Figure 7.3.
APPENDIX B-1 TOWER RELIEF LOAD
Develop an improved methodology to calculate energy balances, using process simulation, for
steam stripped crude towers or absorbers which used high boiling point materials to recover
light ends from gases. Currently, the spreadsheet is best suited for simple fractionation towers
which separate relatively close boiling point materials.
APPENDIX B-2 REACTOR LOOP RELIEF LOAD
Develop a spreadsheet which can calculate change of energy balances in a reactor loop. This
spreadsheet will include simplified exchanger rating methods for different kinds of exchangers.
Chap0-r1.doc
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REV 1 LIST
The following summarizes the changes made in the Revision 1 of the Flare System Manual for
all items other than cosmetic items:
Cover Page:
Notes, contact name, date.
Chapter 0:
Added items to Sections 3.11.5.4 & 7.5.1.3.
Delete item 1 under Appendix B-1.
Chapter 1:
Edited Item 5 in Table 1.2.
The following items were modified in Table 1.1:
• Edited Item 1 Under Task, Secondary/Consultant Responsibility
and Notes
• Edited Item 2. Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 2.D Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 3.B Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 3.C Under Primary Responsibility,
Secondary/Consultant Responsibility
• Added Item 3.D
• Edited Item 4.B Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Added Item 4.D
• Edited Item 5 Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 6 Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 7 Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 8 Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 9 Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 10 Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
• Edited Item 11 Under Primary Responsibility,
Secondary/Consultant Responsibility and Notes
Edited Figure 1.2 to reflect Table 1.1 editions
Chapter 2:
Added English units.
Edited Section 2.1.5.1, item 2 for liquid services.
Chapter 3:
Added English units.
Added equations for English and Metric (Bar) units to Equations
3.1,3.4,3.6,3.7,3.8).
Equation 3.8 was also corrected for a typo.
Rev1 List.doc
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REV 1 LIST
Edited Section 3.2.4 to clarify when to specify Liquid Trim Valves.
Edited Section 3.11.5.4 to include discussion on HEM method for two
phase flow.
Figure 3.1 was replaced with the correct picture.
Edited the constant of Equation 3.6 (metric) from 1.342 to 1.000 in
Section 3.11.1.5.
Chapter 4:
Added English units.
Added equations for English and Metric (Bar) units to Equations
4.1,4.3a,4.3b,4.5,4.6a,4.6b,4.11,4.12,4.13,4.21).
Section 4.3.1.1 was edited.
Heater Duty sub-section under section 4.3.3.2 was edited
Chapter 7:
Section 7.5.1.3.e was edited.
Table 7.1, item 8 was edited. Item 9 was added.
Figure 7.3, Note 2 was corrected.
Appendix B-1:
Rewritten with new examples and backup calculations.
Appendix B-3:
Rewritten with new example.
Appendix B-4:
Rewritten with new example.
Appendix B-5:
More explanation on examples.
Appendix D-1:
New copies of Excel spreadsheets, new examples, and reference of
location of the electronic spreadsheets.
Rev1 List.doc
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TABLE OF CONTENTS
1.0
INTRODUCTION
1.1
DESIGN CONSIDERATIONS
1.2
1.1.1
Precedence of Laws, Local Regulations, Client Standards,
Design Codes, Client Guides and this Manual
1.1.2
Design Objectives
1.1.3
Design Impact Factors
1.1.4
Administrative Procedures
DESIGN RESPONSIBILITIES
1.2.1
Establish Design Philosophy and Standards
1.2.2
System Assessment
1.2.3
Relief Source Identification
1.2.4
Preliminary PSV and Vessel Nozzle Sizing
1.2.5
Final Data Sheet Preparation
1.2.6
Final Relief Load Computation
1.2.7
“As Purchased” Equipment Performance Review
1.2.8
Relief Device Installation Review
1.2.9
Monitor Design Changes
1.2.10 Engineering Documentation
1.3
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CODES, STANDARDS AND PRACTICES
1.3.1
ASME Boiler and Pressure Vessel Code
1.3.2
API Publications
1.3.3
NFPA Standards
1.3.4
ANSI Standards
1.3.5
International Conference of Building Officials (ICBO)
1.3.6
American Institute of Steel constructors (AISC)
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1.4
1.3.7
American Society for Testing Materials (ASTM)
1.3.8
American Welding Society (AWS)
DESIGN GUIDE SUMMARY
1.4.1
Establish Design Pressure of Vessels and Piping (Chapter 2)
1.4.2
Establish Design Temperature of Vessels and Piping (Chapter 2)
1.4.3
Select Type of Relieving Device (Chapter 3)
1.4.4
Establish Individual Relief Loads (Chapter 4)
1.4.5
Calculate Required Relief Device Orifice Area (Chapter 3)
1.4.6
Review Disposal Options (Chapter 6)
1.4.7
Establish Equipment Depressuring Requirements (Chapter 4)
1.4.8
Size Thermal Relief Valves (Chapter 4)
1.4.9
Evaluate Process Flow Loops (Chapter 4)
1.4.10 Evaluate Total Relief Loads to the Flare, by Contingency,
to Include Depressuring (Chapter 5)
1.4.11 Consider Mitigation for Relief Load Reduction (Chapter 5)
1.4.12 Review Depressuring Loads for Time Smoothing (Chapter 5)
1.4.13 Review and Perhaps Modify Control Valves
for Favorable Control Actions (Chapter 5)
1.4.14 Size Relief Valve Piping Inlet/Outlet (Chapter 7)
1.4.15 Establish Required Purging Rates by Converting
Velocities Given Below into Rates [lb/hr (kg/hr) or SCFH (nm3/hr)]
for the Relief Piping (Chapter 7)
1.4.16 Select and Specify the Following Equipment
where Appropriate (Chapter 8)
1.4.17 Develop Flare Stack and Tip Details (Chapter 9)
Table 1.1
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Pressure Relief System Design Responsibilities
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TABLE OF CONTENTS
Table 1.2
Essential Criteria for Flare and Relief System
Figure 1.1 Typical Relief System Engineering Schedule
Figure 1.2 Typical Relief System Activity Flow Chart
2.0
DESIGN PRESSURE AND TEMPERATURE SELECTION
2.1
DESIGN PRESSURE SELECTION
2.2
2.3
Index-r1.doc
2.1.1
Operating Pressure
2.1.2
Maximum Operating Pressure
2.1.3
Settling Out Pressure
2.1.4
Design Pressure
2.1.5
Design Pressure Selection
2.1.6
Design Vacuum
2.1.7
Maximum Allowable Working Pressure (MAWP)
DESIGN TEMPERATURE SELECTION
2.2.1
Definition of Operating Temperature
2.2.2
Maximum Operating Temperature
2.2.3
Definition of Design Temperature
2.2.4
Design Temperature Selection
PSV RELATED PRESSURES
2.3.1
PSV Set Pressure for Vessels
2.3.2
Spring Setting (Cold Differential Test Pressure)
2.3.3
Permissible Overpressure or Accumulation
2.3.4
Superimposed Back Pressure
2.3.5
Built-Up Back Pressure
2.3.6
Back Pressure
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2.4
2.3.7
Pressure Tolerances
2.3.8
Blowdown Pressure
EQUIPMENT RERATING
Figure 2.1 Typical Pressure Levels per API RP 521
Figure 2.2 Allowable Design Stress vs. Temperature
3.0
RELIEVING DEVICES
3.1
INTRODUCTION
3.2
TYPES OF PRESSURE RELIEF DEVICES
3.3
3.4
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3.2.1
Safety Valves
3.2.2
Relief Valves
3.2.3
Safety-Relief Valves
3.2.4
Liquid Trim Relief Valves
3.2.5
Pilot Operated Pressure Relief Valves
3.2.6
Rupture Disks
3.2.7
Non-ASME Devices
CODES AND STANDARDS
3.3.1
ASME Section I
3.3.2
ASME Section VIII
3.3.3
ANSI/API Standard 526
3.3.4
API RP-520, Part I
3.3.5
Testing and Certification
3.3.6
Code Stamps
CONVENTIONAL PRESSURE RELIEF VALVES
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3.5
3.6
3.4.1
Operating Characteristics
3.4.2
Applications
3.4.3
Design Considerations
BALANCED BELLOWS PRESSURE RELIEF VALVES
3.5.1
Operating Characteristics
3.5.2
Applications
3.5.3
Design Considerations
LIQUID TRIM RELIEF VALVES
3.6.1
Operating Characteristics
3.6.2
Applications
3.6.3
Design Considerations
3.7
SPECIAL FEATURES
3.8
PILOT OPERATED PRESSURE RELIEF VALVES
3.9
3.10
Index-r1.doc
3.8.1
Operating Characteristics
3.8.2
Applications
3.8.3
Design Considerations
3.8.4
Special Features
RUPTURE DISKS
3.9.1
Operating Characteristics
3.9.2
Applications
3.9.3
Design Considerations
3.9.4
Rupture Disk Burst Pressure Example
3.9.5
Special Features
OTHER TYPES OF PRESSURE RELIEF DEVICES
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3.10.1 Surface Condenser Pressure Relief Valves
3.10.2 Sentinel Valves
3.10.3 Pressure/Vacuum Breather Valves
3.10.4 Explosion Hatches
3.10.5 Non-ASME Pressure Relief Valves
3.10.6 Liquid Seals
3.10.7 Vacuum Relief Valves
3.11
PRESSURE RELIEF DEVICE SIZING
3.11.1 API Sizing Equations
3.11.2 Manufacturer’s Equations
3.11.3 Pilot Operated Pressure Relief Valves
3.11.4 Safety Valves
3.11.5 Sizing Procedures
3.11.6 Rupture Disk Sizing
3.12
REFERENCES
Table 3.1
API Nozzle Sizes and Areas
Table 3.2
Effective and Actual Areas/Coefficients of Discharge
Figure 3.1
Cross Section of Conventional Pressure Relief Valve
Figure 3.2-A Operating Characteristics of Conventional
Safety Relief Valves in Vapor Service
Figure 3.2-B Operating Characteristics of Conventional Spring
Opposed Pressure Relief Valve in Liquid Service
Figure 3.2-C Operating Characteristics of Liquid Trim Pressure
Relief Valve in Liquid Service
Index-r1.doc
Figure 3.3
Cross Section of Balanced Bellows Pressure Relief Valve
Figure 3.4
Cross Section of Piston Type Pilot Operated Relief Valve
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Figure 3.5
Cross Section of Diaphragm Type Relief Valve
Figure 3.6
Conventional Tension Loaded Rupture Disk
Figure 3.7
Prescored Tension Loaded Rupture Disk
Figure 3.8
Composite Disk
Figure 3.9
Reverse Buckling Disk with Knifes
Figure 3.10
Prescored Reverse Buckling Disk
Figure 3.11
Graphite Disk
Figure 3.12
Rupture Disk Telltale Installation
Figure 3.13
KB Versus Back Pressure for Conventional Pressure
Relief Valves
Figure 3.14
Back Pressure Sizing Factor KB for Balanced Bellows
Pressure Relief Valve (Vapors and Gases)
Figure 3.15
Typical Back Pressure Correction Factor, KW , for
Liquid Service Balanced Bellows Valve (Vapor or
Liquid Trim)
Figure 3.16
Typical Overpressure Correction Factor, KP, for
Conventional Pressure Relief Valve in Liquid Service
Figure 3.17
Rupture Disk Burst Pressure and Manufacturing
Range Tolerances
4.0
DETERMINATION OF INDIVIDUAL RELIEF LOADS
4.1
BASIC PHILOSOPHY
Index-r1.doc
4.1.1
Process Evaluation Basis
4.1.2
Double Jeopardy
4.1.3
Utility Losses
4.1.4
Unsteady State Conditions
4.1.5
Block Valves, Check Valves and Control Valves
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4.1.6
Control System Response
4.1.7
Operator Intervention
4.1.8
Heat Transfer Equipment Performance
4.1.8.1 Air Cooled Exchangers
4.1.8.2 Shell and Tube Exchangers
4.1.8.3 Fired Heaters
4.1.9
4.2
Use of DIERS Methodology
CAUSES OF OVERPRESSURE
4.2.1
General
4.2.2
Operator Error
4.2.3
Utility Failure
4.2.4
Local Equipment/Operation Failure
4.2.5
External Fire
4.2.6
Depressurization
4.2.7
Thermal Expansion
Table 4.1
4.2.8
Chemical Reaction
4.2.9
Miscellaneous
Table 4.2
4.3
Cubical Expansion Coefficient
Bases for Relief Capacities under Selected Conditions
FRACTIONATION AND DISTILLATION
4.3.1
System Description
4.3.2
Causes of Overpressure
4.3.3
Heat and Material Balance Considerations
4.3.3.1 Basic Assumptions for Relief Case
Heat and Material Balance
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4.3.3.2 Heat Balance for Upset Conditions
4.4
4.3.4
Maximum Capacity
4.3.5
Determination of Relief Loads
REACTOR LOOPS
4.4.1
System Description
4.4.1.1 Process Flow
4.4.1.2. Start-of-Run and End-of-Run Conditions
4.4.1.3 Reaction Process Characteristics
4.4.1.4 Alternate Operation Modes
4.4.2
Causes of Overpressure
4.4.3
Heat and Material Balance Considerations
4.4.3.1 Basic Assumptions for Operational Upsets
4.4.3.2 Reactor Yields
4.4.3.3 Condensation Curves
4.4.4
Pressure Profiles
4.4.4.1 Operating Pressure Profiles
4.4.4.2 Design Pressure Profile
4.4.4.3 Relieving Pressure Profile
4.4.4.4 Settle-Out Pressure
4.4.5
Pressure Relief and Depressuring Facilities
4.4.5.1 Code Criteria
4.4.5.2 Location of Pressure Relief Valve
4.4.5.3 Presence of Block Valves in the Loop
4.4.5.4 Pilot Operated Pressure Relief Valve Applications
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4.4.5.5 Depressuring
4.4.6
Maximum Capacity
4.4.7
Determination of Relief Loads
4.4.7.1
Loss of Feed
4.4.7.2
Loss of Effluent Cooling
4.4.7.3
Loss of Quench
4.4.7.4
Recycle Compressor Failure
4.4.7.5
Utility Failure
4.4.7.6
Control Failure
4.4.7.7
Blocked Exits
4.4.7.8
Abnormal Heat Input
4.4.7.9
Change in Feed Composition
4.4.7.10 Chemical Reaction
4.4.7.11 Fire
4.5
4.6
LIQUID FILLED SYSTEMS
4.5.1
Blocked Discharge
4.5.2
Thermal Relief
MECHANICAL EQUIPMENT
4.6.1
Pumps
4.6.2
Compressors
4.6.3
Mechanical Driver Considerations
Table 4.3
4.7
HEAT EXCHANGER TUBE RUPTURE
4.7.1
Index-r1.doc
Condensing Turbines. Atmospheric Safety Valves Sizes
Determining Required Relief Flow Rate
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4.8
4.7.2
Steady State Relief Analysis
4.7.3
Dynamic Relief Analysis
4.7.4
Relief Devices and Locations
4.7.5
Double Pipe Exchangers
FIRE
4.8.1
Basis Assumptions for Fire Case Relief Analysis
4.8.2
Heat Flux Equations
4.8.3
Determination of Relief Loads for Equipment Containing Liquid
4.8.4
Relief Loads for Vessels Containing Vapor
4.8.5
Depressuring
4.9
CHEMICAL REACTIONS
4.10
ATMOSPHERIC STORAGE TANK PROTECTION
4.10.1 Relief Device Accumulation
4.10.2 Non-refrigerated Aboveground Tanks
4.10.3 Refrigerated Aboveground and Belowground
4.10.4 Means of Venting
4.11
REFERENCES
Figure 4.1 Isothermal Flow of Compressible Fluids Through
Pipes at High Pressure Drops
5.0
OVERALL RELIEF SYSTEM LOAD EVALUATIONS
5.1
INTRODUCTION
5.2
FLARE LOAD CALCULATIONS
Index-r1.doc
5.2.1
General Methodology
5.2.2
Determination of Relief System Loads
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5.3
5.4
5.2.2.1
Determination of Area Fire Loads
5.2.2.2
Utility Failure
5.2.2.3
Other Contingencies
5.2.2.4
Loads from Depressuring Systems
FLARE LOAD MINIMIZATION
5.3.1
Background
5.3.2
System Design and Modifications
5.3.3
General Approaches
5.3.4
Risk Based Assessment
5.3.5
Recommendations for Relief Load Mitigation Instrumentation
Table 5.1
Pump Autostart Load Reduction Credits
Table 5.2
Dual Loop Shutdown System Load Reduction Credits
5.3.6
Dynamic Simulation
5.3.7
Probability Analysis
REFERENCES
Figure 5.1 Triple Loop Shutdown System
6.0
RELIEF MATERIAL RECOVERY AND DISPOSAL
6.1
GENERAL
6.2
DISPOSAL OPTIONS
Index-r1.doc
6.2.1
Discharge to Atmosphere
6.2.2
Discharge to Grade or Sewer
6.2.3
Discharge to a Process Vessel
6.2.4
Discharge to a Closed Collection System
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6.3
HAZARD AND RISK ASSESSMENT
6.4
ENVIRONMENTAL FACTORS
6.5
VAPOR RELEASE CRITERIA
6.6
6.7
6.8
Index-r1.doc
6.5.1
General
6.5.2
Atmospheric Release Criteria
6.5.3
Safety Review
LIQUID RELEASE CRITERIA
6.6.1
Non-Hazardous Streams
6.6.2
Non-Hazardous Hydrocarbons
6.6.3
Hazardous Streams
6.6.4
Two Phase Releases
6.6.5
Prevention of Liquid Releases
6.6.6
Pressure Relief Device Failure
DISPOSAL INTO A PROCESS
6.7.1
Capacity
6.7.2
Destination Pressure
6.7.3
Process Upsets
6.7.4
In-Service Requirements
CLOSED DISPOSAL SYSTEMS
6.8.1
Intermediate Collection Systems
6.8.2
Flare Systems
6.8.3
Vapor Recovery
6.8.4
Incinerators & Burn Pits
6.8.5
Liquid Handling Systems
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6.8.6
6.9
6.10
Treating Systems
DESIGN CONSIDERATIONS
6.9.1
Atmospheric Releases
6.9.2
Intermediate Collection Systems
6.9.3
Flare Systems
6.9.4
Vapor Recovery
6.9.5
Incinerators
6.9.6
Liquid Handling Systems
6.9.7
Treating Systems
REFERENCES
Table 6.1
Typical Threshold Limit Values for Toxic or
Hazardous Chemicals Found in Refineries
Figure 6.1 Typical Flare Gas Recovery System
Figure 6.2 Flare Gas Recovery Inlet Pressure Control System
Figure 6.3 Typical Quench Drum
Figure 6.4 Typical Scrubber System
7.0
RELIEF SYSTEM PIPING & SEALING/PURGING
7.1
DESIGN CONSIDERATIONS
Index-r1.doc
7.1.1
Piping Layout Guidelines
7.1.2
Design Temperature
7.1.3
Design Pressure
7.1.4
Stress
7.1.5
Isolation Valves
7.1.6
Design Criteria for Relief Valve Inlet Piping
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7.2
7.3
7.4
7.5
7.6
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7.1.7
Design Criteria for Relief Headers
7.1.8
Piping Metallurgy
7.1.9
Winterization, Safety Insulation and Steam Tracing
LINE SIZING
7.2.1
Relief Valve Inlet/Outlet Piping Sizing
7.2.2
Line Sizing of the Main Relief Header
COMPUTER MODELING OF FLARE HEADERS
7.3.1
Network Method
7.3.2
Flare Method
7.3.3
Contingency Allowance
7.3.4
Pipe Roughness ( ε )
7.3.5
Hydraulic Evaluation
FLOW METERING
7.4.1
Design Discussion
7.4.2
Methods
SEALING AND PURGING
7.5.1
Sealing
7.5.2
Purge Gas
REFERENCES
Table 7.1
PSV Inlet/Outlet Calculations, Design Criteria
Table 7.2
Orifice/Inlet Area Ratio for Standard Relief Valves
Table 7.3
Maximum Allowable Equivalent Lengths of Inlet
Piping to Comply with 3% Inlet Loss Criteria for Relief Valve
Table 7.4
Orifice/Outlet Area for Standard Relief Valves
Table 7.5
Typical Outlet Nozzle Lengths
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Table 7.6
Typical Reducer Angles
Table 7.7
Typical Weld Neck Flange Lengths
Table 7.8
Typical Friction Factors for Clean Carbon Steel Pipe
Figure 7.1 Baffle Type Seal
Figure 7.2 Labyrinth Type Seal
Figure 7.3 Vertical Water Seal Drum
Figure 7.4 Flare Purge Gas Supply
Figure 7.5 Typical Pressure Relief Valve Installation:
Atmospheric (Open) Discharge
Figure 7.6 Typical Pressure Relief Valve Installation:
Closed System Discharge
Figure 7.7 Typical Pressure Relief Valve Mounted on
Process Line
Figure 7.8 Typical Pressure Relief Valve Mounted on
Long Inlet Pipe
Figure 7.9 Typical Pilot-Operated Pressure Relief
Valve Installation
Figure 7.10 Typical Rupture Disk Assembly Installed in
Combination with a Pressure Relief Valve
Figure 7.11 Typical Pressure Relief Valve Installation
with an Isolation Valve
8.0
KNOCKOUT, BLOWDOWN, SEAL, QUENCH DRUMS AND PUMPS
8.1
KNOCKOUT DRUM
8.2
Index-r1.doc
8.1.1
Purpose
8.1.2
Design Parameters
8.1.3
Design Details
BLOWDOWN DRUM
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8.3
8.4
8.2.1
Purpose
8.2.2
Design Parameters
8.2.3
Design Details
SEAL DRUM
8.3.1
Purpose
8.3.2
Design Parameters
8.3.3
Design Details
QUENCH DRUMS
8.4.1
Purpose
8.4.2
Design Parameters
8.4.3
Design Details
8.5
PUMPS
8.6
REFERENCES
Table 8.1
Table of Geometry for Circles and Arcs
Figure 8.1 Typical Horizontal Knockout Drum
Figure 8.2 Drag Coefficient, C
Figure 8.3 Typical Horizontal Blowdown Drum
Figure 8.4 Typical Horizontal Seal Drum
Figure 8.5 Schematic for Combined Ground Flare and
Elevated Flare
Figure 8.6 Typical Operating and Emergency Flares
Figure 8.7 Typical Quench Drum (Condensable)
Figure 8.8 Typical Quench Drum (Emergencies)
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9.0
FLARE
9.1
DESIGN DISCUSSION
9.2
9.3
9.4
9.5
9.6
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9.1.1
Selection of Flare Stack Location
9.1.2
Flare System Data Sheet
TYPES OF FLARES
9.2.1
Discussion
9.2.2
Elevated Flare
9.2.3
Ground Flares
9.2.4
Offshore Platform Flares
FLARE SYSTEM METALLURGY
9.3.1
Hydrocarbon Flaring
9.3.2
H2S Flaring
ELEVATED FLARE SIZING
9.4.1
Discussion of Sizing Methods
9.4.2
Nomenclature
9.4.3
Stack Diameter
9.4.4
Stack Height
9.4.5
Radiation Considerations
9.4.6
Equipment Surface Temperature
GROUND FLARE SIZING
9.5.1
Enclosed Ground Flares
9.5.2
Open Pit Ground Flares
9.5.3
Burn Pit
SMOKELESS FLARING
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9.7
9.8
9.9
9.10
9.6.1
Smokeless Flaring Requirements
9.6.2
Steam Injection
9.6.3
Air Assisted Flaring
9.6.4
Miscellaneous
9.6.5
Smokeless Flaring Control
FLARE TIP DESIGN OPTIONS
9.7.1
Flare Tip Characteristics
9.7.2
Open Pipe Flare Tip
9.7.3
Forced Draft (Air Assisted) Flare Tips
9.7.4
Multi Tip Flares
9.7.5
Coanda Flare Tip
9.7.6
High Velocity Tips
NOISE & ENVIRONMENTAL
9.8.1
Noise Standards
9.8.2
Noise Discussion
9.8.3
Environmental
FLARE IGNITION
9.9.1
Discussion
9.9.2
Pressure Ignitor
9.9.3
Electronic Ignitor
9.9.4
Atmospheric Ignitor
9.9.5
Pilot Monitoring
REFERENCES
Table 9.1
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Comparison of Flare Types
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Table 9.2A Lower Limits of Flammability of Gases and Vapors in Air
Table 9.2B Lower Limits of Flammability of Gases and Vapors in Air
Table 9.3
Recommended Surface Emissivity Values (εs)
Table 9.4
Air Required for Stoichiometric Combustion of Gases
Figure 9.1 Stack & Flame Geometry
Figure 9.2 XL versus SL
Figure 9.3 Temperature of Steel vs Time of Exposure
Figure 9.4 Typical Enclosed Ground Flare Configuration
Figure 9.5 Typical Burn Pit
Figure 9.6 Typical Air-Assisted Flare System
Figure 9.7 Steam/Hydrocarbon Ration vs Flare Gas Molecular
Weight for Smokeless Flaring
Figure 9.8 Conventional Pipe Flare
Figure 9.9 Conventional Flare with Steam Water Spray
Figure 9.10 High Velocity Tip
Figure 9.11 Air Assisted Flare Tip (Top View)
Figure 9.12 Coanda Nozzle (Internal)
Figure 9.13 Coanda Flare (External)
Figure 9.14 Offshore Flare Support Types
APPENDIX A - NOMENCLATURE
APPENDIX B - SAMPLE CALCULATIONS
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B-1
Tower Relief Load
B-2
Reactor Loop Relief Load
B-3
Rupture Tube Relief Load
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B-4
Fire Relief Load
B-5
Rating of the Horizontal Flare K.O. Drum
B-6
Flare Radiation
B-7
Dynamic Simulation
B-8
Flare System Hydraulics Calculations
APPENDIX C – PSV SIZING SOFTWARE
APPENDIX D - Flare System Calculation Spreadsheets
D-1
D-2
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Flare System Spreadsheet Calculations
Typical Calculation Index
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1.0
INTRODUCTION
1.1
DESIGN CONSIDERATIONS
1.1.1
Precedence of Laws, Local Regulations, client standards, Design Codes, client
guides and this Manual
It is intended that this manual supplement rather than replace or supersede any of
the laws, regulations, standards, design codes or guides listed in Section 1.3.0. A
thorough knowledge of these design criteria is essential for a safe design. Any
apparent conflicts with this manual are to be resolved in such a manner as to satisfy
the order of the following precedence:
1. National Laws
2. Local Regulations
3. Client Standards
4. Design Codes
5. Client Design Guides
6. Fluor Daniel Flare System Manual
7. Industry Standards and Guidelines
1.1.2
Design Objectives
The purpose of a flare system is to safely limit the pressure on operating equipment
and interconnecting piping to the maximum allowable pressure. The relieving
system size is dictated by the volume to be relieved and the pressure available to
transfer this volume to the flare.
This manual achieves this purpose and more when the design objectives listed
below are met:
1.1.3
•
The system provides adequate safety for personnel and equipment, thus
concurring with all safety laws, design codes and standards.
•
Atmospheric emissions are lower, enhancing the environmental acceptability
of the plant.
•
Energy is conserved by the recovery and reuse of valuable hydrocarbons as
fuel, providing added profits to the client.
•
Plant siting problems are minimized by the reduction of flare emissions,
luminescence, noise and smoke.
Design Impact Factors
Relief and flare system designs begin with a collection of preliminary engineering
information which may impact on the proposed design. A partial list of these would
include:
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Heat and material balances of process units.
•
Power distribution oneline diagram (conceptual).
•
The anticipated relieving quantities for various emergency conditions such as
cooling water failure, total power failure, etc., where the plant has been
designed in a traditional conservative manner.
•
The anticipated relieving quantities for the same emergency conditions as in
the previous item for a plant using improvements which might incorporate
some, or all, of the following functions to minimize flare quantities and
contain, retain, and recirculate hydrocarbons which would otherwise be
vented to the atmosphere:
°
Design of power distribution system to minimize relieving quantities.
°
Increased design pressures of key equipment.
°
Highly reliable double-lead electrical systems from dual power grids
to ensure high on-stream reliability.
°
Instrumentation to lock out reboiler heat sources for fractionation
towers.
°
Reliable driver selection for reflux pumps in key fractionating
systems.
°
Cascading hot vapor relief streams through compatible cooling
systems to maximize liquid condensation before relieving.
•
Operating and investment costs for the flare and relief system components.
•
The expected frequency of normal operational upsets and of major
emergency situations, which will activate the flare system.
•
The methods used for venting gases generated during start-up, shutdown
and depressuring operations.
•
Recovery of gases from sources previously vented to the atmosphere such
as atmospheric storage tanks, sour water storage tanks, and compressor
distance piece vents.
•
Liquid recovery from the pumpout and/or blowdown header.
•
Quench and scrubbing systems in the flare header to recover valuable
substances.
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1.1.4
•
Segregation of sweet and sour flare gases into two headers to provide an
economical means for recovery, treating, and reuse of valuable vapor and
liquid components.
•
Segregation of high and low pressure headers to minimize required piping
sizes.
•
Equipment required for a Vapor Recovery System to contain, retain and
recirculate the gases and liquids from sources such as relief valve leaks,
minor upsets, and header purge gases.
•
Consideration of the benefits of multiple combustion systems to handle small
and large flare quantities in separate systems such as open pit combustion,
ground flares, and enclosed thermal oxidizers.
•
The environmental and safety standards which must be met in the area
surrounding the process unit protected by the flare system.
•
The concerns of local communities for the impact the plant will have on
them.
Administrative Procedures
Administrative procedures have an important economic role in the safe design and
operation of pressure relief systems. Application of administrative procedures,
however, places a burden on management of the refinery for maintenance of the
required procedures. For this reason, these procedures are to be applied only when
the benefits exceed the burdens.
Administrative procedures which are related to pressure relief systems must be
clearly defined, clearly communicated to unit operators and strictly enforced. Plant
management has direct responsibility for accepting the risks that can be associated
with administrative procedures and for assuring that administrative procedure
policies are established and enforced.
A partial listing of possible administrative procedures follows:
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Lock (or car seal) procedures for block valves associated with pressure relief
valves. The procedures should include a list of all block valves which are
required to be locked in position, definition of who is authorized to unlock and
move block valve positions, procedures for maintaining logs of locked block
valve movements and definition of how the procedures will be enforced.
•
Requirements that equipment be continuously attended during certain
operations, such as when a pressure relief valve is blocked in or when
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equipment is operated in a mode, such as steam out or pump out, that it is
known the pressure relief system is not designed to protect against.
1.2
•
Limitations on modification of equipment without the proper engineering
review of the effect on the pressure relief system. Examples of these types
of limitations are restrictions on changes of pump impeller sizes or turbine
driver speed settings, operating control valves with their bypass valves
partially or fully open, adjustment or removal of control valve minimum or
limit stops or revisions to control valve internal trim.
•
Operating procedures for shutting down a unit under pre-identified failure
conditions.
•
Vent and drain procedures for equipment maintenance.
DESIGN RESPONSIBILITIES
1.2.1
Establish Design Philosophy and Standards
The Unit Process Engineer determines causes of overpressure and flows to the
relief and flare system for the process units. The Upset Condition Checklist and
Individual Relief Summary Sheet are utilized for these tasks. Process data for relief
devices are summarized and transmitted to the Control systems engineer on the
Relief Valve Calculations Sheet.
The Unit Process Engineer for the relief and flare system reviews the relief loads for
the various units and determines the disposal methods for the streams from each
unit and maximum flow to each disposal system and calculates the sizes for all lines
including relief device inlet piping. Other duties include developing process and
system flow diagrams, providing process operating and design data for equipment
and piping, and reviewing piping isometric drawings.
The Flare System Process Engineer provides assistance in development of the
design philosophy and assures that the client standards and guides and applicable
Laws and Codes are properly applied to the design. He will review plot plans for
location of equipment containing volatile liquids and participate in the system design
development and reviews as the project proceeds.
Design and maintenance of pressure relief systems is a multi-discipline activity
which begins with the conceptual design of a new process unit, utility or offsite
facility and continues throughout its use. Once a pressure relieving system has
been designed and installed, legal obligations and responsible practice require that
records of the engineering design basis, inspection and maintenance history for the
system be kept.
The sizing basis for pressure relief systems is often directly related to process
variables such as flow rates, equipment capacities or operating characteristics,
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process fluid properties, etc. As a consequence, the design requirements for
pressure relief systems may change if modifications are made to system equipment
or the piping configuration or if the process equipment is operated differently than
planned for in the original design. The sizing basis for a pressure relief system must
be reviewed every time a significant change in minimum or maximum process rate,
stream composition, operating conditions, equipment capacity or equipment line-up
is made.
The various tasks and responsibilities associated with design and ownership of
pressure relief systems are discussed in this Section. The groups typically
responsible for each task are identified for the purposes of discussion only and
actual assignments should be designated by management.
Figure 1.1 and 1.2 diagram the typical activities in designing a pressure relief
system. Typical tasks and responsibilities are listed in Table 1.1. Finally, a Flare
and Relief System Checklist, Table 1.2, is included at the end of Section 1. This
table is not intended to be all inclusive with respect to flare system design but rather
to assure that the more important design parameters have been addressed in the
design.
1.2.2
System Assessment
1.2.2.1
Set Equipment Design Conditions
As part of process design development, the Unit Process Engineer
determines the maximum pressure and temperature for which each piece
of equipment must be designed. Section 2 of this manual discusses the
selection of design pressures and temperatures. This activity should
recognize the interaction between equipment design and overpressure
protection.
In addition to consideration of normal and maximum operating pressure
and temperature, the design condition selected should include
consideration of:
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Pressure and temperature excursions due to process upsets
•
Maximum pressure of external sources
•
Maximum pressures that rotating equipment can develop under
both normal and abnormal conditions
•
Performance limitations of pressure relief devices
•
Potential relief stream disposal problems.
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The need for pressure relief systems cannot be eliminated by
specification of higher design conditions. Often however, pressure relief
system designs can be simplified or minimized by careful selection of
mechanical design conditions.
1.2.2.2
Review Plot Plan for Fire Exposure
The plot distribution of vessels and exchangers containing volatile liquids
must be reviewed. Too close a grouping of these equipment types could
lead to a very large relief during a fire, resulting in an excessively large
relief header. This activity must be performed by the Flare System
Process Engineer early in the design to prevent expensive and untimely
changes to the plot plan later in the design phases.
1.2.2.3
Relief Stream Disposal
Relief streams must be disposed of in a safe, economical and
environmentally acceptable manner.
As part of the process design development, the methods of disposal of
relief streams must be determined by the Unit Process Engineer. The
evaluation for this decision must consider the following:
•
Local environmental requirements: Unit operating permits may
be contingent upon disposal of all or some relief streams to a
closed disposal system. Releases of some types of streams to
atmosphere may have to be reported to pollution control
authorities.
•
Potentially dangerous or toxic relief streams: In refineries,
relatively few relief streams are considered to be toxic. However,
release of flammable liquid or two-phase mixtures, or of high
molecular weight condensable vapors, may pose unacceptable
hazards and must be avoided or contained.
Guidelines for disposal of relief streams are discussed in Section 10.
1.2.2.4
System Review
Before a process design is released for detailed engineering, the Unit
Process Engineer, Flare System Process Engineer, Control Systems
Engineer, Mechanical Engineer, Piping Engineer and Refinery
Operations should conduct a joint review under the direction of the
Project Engineer of the preliminary pressure relief system design. The
purpose of this review is to determine where unit economics or safety
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might be improved by upgrading equipment design conditions. Typically
this review is performed as part of the P&ID review.
1.2.3
Relief Source Identification
The services which require pressure relief devices and their physical location are
provisionally established during process design and confirmed during the detailed
design phase. The basic requirements for determining where pressure relief
devices need to be located are discussed in Section 7.
Final determination of pressure relief device services and locations is dependent
upon completion of the P&ID and the piping isometrics. However, the need for and
location of the major pressure relief devices must be identified during process
design by the Unit Process Engineer. Be aware that during detailed design, “aspurchased” equipment performance characteristics or addition of equipment, piping,
block valves, or control valves may change the pressure relief system requirements.
Each project should conduct a pressure relief system safety review at the
appropriate stage of design development, as discussed in Section 1.2.8.
1.2.4
Preliminary PSV and Vessel Nozzle Sizing
1.2.4.1
Preliminary Relief Load Calculations
When the process design is complete and sufficient mechanical
information is available, the Unit Process Engineer is responsible for
determination of relieving rates and completion of the Relief Valve
Calculation Sheet for each service and a unit summary on the Individual
Relief Load Summary.
The basis and techniques for developing relief flow rates are discussed
in Section 4.
Evaluated relief cases and calculated relief loads for all major services,
such as fractionation towers or reactor loops, should be reviewed at this
time with control systems engineering. Other pressure relief services
which are deemed critical either due to their complexity or the magnitude
or quantity of their relief loads should also be included in this preliminary
review.
1.2.4.2
Preliminary Valve Sizing
The initial sizing of pressure relief devices by the Control Systems
Engineer should begin as soon as there is enough information for an
effective analysis. This is usually before final mechanical design data
such as rotating equipment performance curves are available. In order
to generate a sizing basis, the unit Process Engineer will have to rely on
reasonable assumptions where firm design information is lacking. It is
the Unit Process Engineer’s further responsibility to monitor purchased
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equipment performance data as well as flow sheet developments
throughout the job to ensure that previously specified relief quantities
remain valid or are modified accordingly.
1.2.4.3
Preliminary Nozzle Sizing
The Unit Process Engineer is responsible for determining the size and
location of nozzles for mounting pressure relief devices on process
vessels. Although final pressure relief device sizing information may not
be available when the process vessel specifications are first issued, all
known pressure relief connections should be indicated on the vessel
sketch, with sizes placed on “hold”, if necessary. When the Control
Systems Engineer receives the Relief Valve Calculation Sheet and
carries out preliminary pressure relief device sizing and selection, the
process engineer can then confirm nozzle sizing following appropriate
hydraulic calculations.
Nozzles for pressure relief valves should be generously sized to minimize
the chance that a change will be required after vessel fabrication has
started. Any requirements such as rounded entrance nozzles should be
specified at this time.
1.2.5
Final Data Sheet Preparation
After the Unit Process Engineer issues the Relief Valve Calculation Form (Section
1.2.4), the Control Systems Engineer performs preliminary sizing calculations and
makes a preliminary selection of pressure relief device types and sizes. Once the
P&ID has been fully developed and all equipment specified, the Unit Process
Engineer finalizes relief loads. At this time, the Control Systems Engineer performs
final sizing calculations and prepares detailed purchase specifications which include
specification of materials of construction, accessories and required code stamps in
addition to the pressure relief device size and type.
Guidelines for pressure relief device sizing and selection are discussed in Section 3.
1.2.6
Final Relief Load Computation
The Unit Process Engineer is responsible for summarizing the case loads for all the
events causing relief to the flare headers and establishing the controlling cases for
design. Some recycle of work occurs at this time to optimize the flare system
design. For this reason, an early start to this activity is important.
1.2.7
“As Purchased” Equipment Performance Review
When the equipment is committed for purchase, the assumptions taken during early
development of the flare and relief system must be confirmed. The Unit Process
Engineer is responsible for performing detailed relief case evaluations based on the
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actual purchased capacities of the equipment such as pumps, compressors, control
valves, etc..
1.2.8
Relief Device Installation Review
Pressure relief devices may operate marginally or not at all if improperly installed.
Additionally, codes contain specific installation requirements and limitations. During
P&ID development, the Control Systems Engineer shall be responsible for reviewing
the designs and advising the Project Engineer where the client standards and local
codes are not being met.
Installation requirements are discussed in Section 7.
As soon as piping sketches are available, the Piping Designer and the Unit Process
Engineer must verify that the pressure relief valve inlet and outlet piping pressure
losses are within acceptable tolerances for the installation. This evaluation will serve
as a check of the inlet loss determination which was done (possibly using an
assumed inlet piping configuration) to establish the vessel relief valve nozzle sizes.
If any pipe or nozzle sizes appear marginal, they should be increased at this time to
avoid costly rework. Evaluation of any services which discharge to a common
blowdown header may have to be deferred until all pressure relief valves connected
to that header have been specified.
Maximum allowable losses and methods for computing them are covered in
Section 7.
During detailed piping design, the Piping Designer is responsible for ensuring that
the piping design meets all criteria for pressure relief device installation. This
includes verification that all block valves are properly specified, the piping is
correctly supported, and that it is adequately designed for anticipated thermal
stresses and reactive forces. If necessary, the piping design should be reviewed
with the Control Systems Engineer. In addition, the Piping Designer has a
continuing responsibility for advising the Unit Process Engineer of any significant
changes in piping arrangement, so that hydraulic calculations can be reviewed as
necessary. Copies of all calculations must be retained and included with the
pressure relief device documentation package.
1.2.9
Monitor Design Changes
As the design development for the flare and relief system proceeds, changes
usually occur in the design basis for both the flare and relief system and the process
being protected. This would include establishment of those items for which
assumptions were originally taken to allow development to proceed. The impact of
these changes on the flare and relief system must be constantly monitored by the
Project Engineer to provide the best system design with respect to operation and
cost. The Unit Process Engineer and the Control Systems Engineer are responsible
to keep the Project Engineer informed of these changes and their impacts. An
example of these types of impacts would be the cooling water pump drivers. If an
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early decision had been made that the drivers would all be electrical then total loss
of power would also mean total loss of cooling water. This impact could be
minimized if turbine drivers were mixed with the motor drivers allowing only partial or
no loss of cooling water during total power failure.
1.2.10 Engineering Documentation
As mentioned earlier, design of pressure relief systems is sensitive to processing
conditions, processing rates, equipment size and equipment configuration.
Whenever a change in a unit’s operation or equipment is planned, the effect upon
pressure relief system design must be evaluated. A key factor in being able to
perform these evaluations is availability of engineering records.
Pressure relief system design is complex, and unless detailed engineering records
of the existing systems are kept, it is often prohibitive, both in terms of cost and
time, to reassemble all the process and mechanical data required and to recreate
process relief calculations every time the design basis for a pressure relief system
needs to be reviewed. Therefore, a complete engineering record for each pressure
relief device should be provided to the client, up to the level of detail consistent with
Fluor Daniel’s scope of work on the project. Responsibility for completion of these
files, updated to include as-built conditions, should be clearly defined with the client
when the files are transferred.
All cases in which special operating limits or procedures form part of the relief
system design basis must be clearly identified. Among the items which must be
defined jointly with the client and provided in the design basis are:
•
A block valve lock and seal administrative policy and a list of all block valves
which are covered by it.
•
Specific operating procedures, such as vent and drain procedures which
apply to the pressure relief system.
•
Identification of limitations on equipment operating ranges or modifications
which may be performed.
•
Allowable line-ups in systems with installed spare pressure relief devices.
The client has the responsibility for ensuring that operators have clearly defined
procedures and restrictions relative to pressure relief systems, that all operators are
properly trained in these procedures and restrictions, and that the procedures and
restrictions are carefully observed.
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1.3
CODES, STANDARDS AND PRACTICES
1.3.1
ASME Boiler and Pressure Vessel Code
•
Section I
Power Boiler
•
Section IV
Heating Boilers
•
Section VIII
Pressure Vessels
These codes are for design of pressure vessels and boilers in excess of 15 psig
2
(1.03 barg, 1.05 kg/cm G) internal pressure. This represents the majority of the
vessels in refinery operations. The primary sections of interest in Section I are PG67 to PG-73. The primary sections of interest in Section VIII are UG-15 to UG-136
and Appendices M and 11.
1.3.2
API Publications
Standards:
API Std 526
Flanged Steel Safety-Relief Valves
This standard specifies dimensions of carbon and alloy
steel safety-relief valves.
API Std 620
Recommended Rules for Design and Construction of
Large, Welded, Low-Pressure Storage Tanks
This standard is for design of low pressure storage
tanks less than 15 psig (1.03 barg, 1.05 kg/cm2 G) but
2
greater than 0.28 psig (0.02 barg, 0.02 kg/cm G).
API Std 650
Welded Steel Tanks for Oil Storage
This standard is for design of atmospheric pressure
storage tanks with internal pressures up to 0.28 psig
(0.02 barg, 0.02 kg/cm2 G).
API Std 2000
Venting Atmospheric and Low-Pressure Storage
Tanks (Non-refrigerated and Refrigerated)
This standard covers the specification of relief valves
for vessels and tanks with design pressures less than
15 psig (1.03 barg, 1.05 kg/cm2 G).
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API Std 2510
Design and Construction of LP-Gas (LPG) Installations
at Marine and Pipeline Terminals, Natural Gas
Processing Plants, Refineries, and Tank Farms
Bulletins:
API Bulletin 2521
Use of Pressure Vacuum Vent Valves for Atmospheric
Pressure Tanks to Reduce Evaporation Losses.
Recommended Practices:
API RP 520
Design and Installation of Pressure Relieving Systems
in Refineries - Part I (Design) and Part II (Installation)
This recommended practice has been accepted as the
most authoritative set of rules for sizing and
specification of individual relief devices. As indicated
above, the first part is for design and the second part
is for installation. It is anticipated that this standard will
be used in conjunction with API RP 521 to provide a
consistent design basis for flare and relief system
design.
This practice is intended for use for relief valves to be
installed on vessels and tanks with design pressures
2
of 15 psig (1.03 barg, 1.05 kg/cm G) or greater.
API RP 521
Guide for Pressure Relief and Depressuring Systems
This recommended practice has also been accepted
as an authoritative set of rules. Its application is for
design of a relief system to safely dispose of the
individual relief loads established by designs
conforming to API RP 520.
API RP 576
1.3.3
Inspection of Pressure-Relieving Devices
NFPA Standards
NFPA 30
Flammable and Combustible Liquids Code
Use this standard for non-refinery low pressure
storage, less than 15 psig (1.03 barg, 1.05 kg/cm2 G).
NFPA 58
Standard for Storage and Handling of Liquefied
Petroleum Gases
Use for non-refinery gas plant LPG storage
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1.3.4
ANSI Standards
B9.1
Safety Code for Mechanical Refrigeration
B19.3
Safety Standard for Compressors for Process
Industries
B31.1
Power Piping
B31.2
Industrial Gas and Air Piping
B31.3
Petroleum Refinery Piping
B31.4
Liquid Petroleum Transportation Piping
B31.5
Refrigeration Piping
B31.6
Chemical Process Piping
B31.8
Gas Transmission and Distribution Piping
B95.1
Terminology for Pressure Relief Devices
1.4.9
NEMA Standards
SM21
Multistage Steam Turbines for Mechanical Drive
Service
SM22
Single Stage Steam Turbines for Mechanical Drive
Service
1.3.5 International Conference of Building Officials (ICBO)
Uniform Building Code (UBC)
1.3.6 American Institute of Steel Constructors (AISC)
Manual of Steel Construction
1.3.7 American Society For Testing Materials (ASTM)
A320
Specification for Alloy-Steel Bolting Materials for Low
Temperature Service.
1.3.8 American Welding Society (AWS)
D.1.1
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Structural Welding Code
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D.14.4
1.4
Classification and Application of Welded Joints for
Machinery and Equipment.
DESIGN GUIDE SUMMARY
This Flare System Manual was developed to provide guidance for the engineering effort
involved in the development of safe and efficient flare and relief systems. The tasks
involved in this development include proactively obtaining the inputs of the project
engineering task force at appropriate times as well as performing the design engineering
functions. For this reason a matrix of approaches appear in this Section. Figure 1.1,
Typical Relief System Engineering Schedule, presents the developmental tasks from Table
1.1 on a timeline to clarify the timeliness of tasks number 1 through 10. The same tasks
are presented again in Figure 1.2, Typical Relief System Activity Flow Chart, to give a
perspective to the engineer developing the relief system. Note the recycle of changes in
Figure 1.2. These recycle activities are frustrating but necessary to the most efficient
overall development of a project. Prior knowledge that these recycles will occur and
contingency planning to minimize the impacts will add significantly to the smooth flow of the
project and a timely completion. The temptation to wait until all the necessary inputs are
available and final must be avoided if project schedules are to be met.
The Essential Criteria for Flare and Relief System, Table 1.2, provides a checklist that
should be referred to by the Lead Process Engineer as the tasks in Table 1.1 are
performed. Because of the uniqueness of each project, additional essential criteria should
be identified and developed by the Lead Process Engineer on a project by project basis. If
any of the criteria in Table 1.2 are interpreted to be in conflict with the later Sections of this
design guide, the detailed guidelines in the individual Sections should be given precedence.
The same is true for the summary of technical tasks below.
The following technical tasks must be performed in the development of a flare and relief
system design. Each task is a summary or condensation of the items which are discussed
in further detail in the referenced location in the following Sections. The use of this
summary presupposes that the engineer is knowledgeable about the background
information contained in these Sections and API RP 520 and 521. Where this is not the
case, the appropriate material should be reviewed.
1.4.1
Establish Design Pressure of Vessels and Piping (Section 2)
•
Vessels - Select the highest of:
1.
2.
3.
4.
5.
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2
Operating pressure plus 25 psi (1.8 bar, 1.8 kg/cm )
Operating pressure times 1.1 for vapor and 1.2 for liquid
2
50 psig (3.4 barg, 3.5 kg/cm G) if a PSV on the vessel PSV relieves
to the flare header
2
30 psig (2.1 barg, 2.1 kg/cm G) if the vessel PSV relieves to
atmosphere
2
15 psig (1.03 barg, 1.05 kg/cm G) if the vessel is vented to
atmosphere
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•
Piping - Select the higher of:
1.
2.
3.
4.
1.4.2
Vessel design pressure for vapor
Vessel design pressure plus static head for liquids
Centrifugal pump shutoff pressure
Positive displacement pump, based on relief valve set pressure,
which is typically set at the lower of:
a.
Casing mechanical design pressure
2
b.
Higher of rated pressure plus 25 psi (1.8 bar, 1.8 kg/cm ) or
110% of rated pressure.
Establish Design Temperature of Vessels and Piping (Section 2)
Set the design temperature at the maximum operating temperature coincident with
the design pressure selected from section 1.5.1 above plus a design margin as
shown:
1.4.3
•
Add 50 °F (28 °C) for operating temperatures up to 650 °F (343 °C)
•
Add 25 °F (14 °C) for operating temperatures over 650 °F (343 °C)
Select Type of Relieving Device (Section 3)
•
Conventional Pressure Relief Valve
•
Spring loaded, built-up back pressure tends to reseat the valve beyond about
10% of set pressure.
Balanced Pressure Relief Valve (Piston & Bellows types)
Spring loaded, the effects of built-up back pressure are reduced greatly by
use of piston or bellows to allow up to 30 or 40% of set pressure without
capacity reduction.
•
Pilot Operated Pressure Relief Valve (Piston & Diaphragm types)
Consists of a main valve and an external pilot valve that can be modulating
or “pop” action.
•
Safety Valve
Primarily for ASME Section I relief service. These spring loaded valves
provide full opening with minimum overpressure.
•
Relief Valve
Spring loaded pressure relief device for liquid relief service.
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•
Safety Relief Valve
Liquid and vapor service. The majority of the refinery relief devices are of
this type. This type valve performs like a safety valve in vapor service (pop
action) and a relief valve in liquid service.
•
Rupture Disk
Rupture disks are used only infrequently for special applications
1.4.4
1.4.5
Establish Individual Relief Loads (Section 4)
•
Establish Sizing Basis for the PSV
•
Add the relief load to the total relief summary on a case by case basis
Calculate Required Relief Device Orifice Area (Section 3)
API Sizing Equations
•
Vapor in Critical Flow
A = [ W / (C Kd P1 Kb) ] [ (T Z) / MW ]0.5
0.5
A = 1.316 [ W / (C Kd P1 Kb) ] [ (T Z) / MW ]
0.5
A = 1.342 [ W / (C Kd P1 Kb) ] [ (T Z) / MW ]
•
a.
Check if Back Pressure Corrections are required by valve
manufacturer (Kb)
b.
Check if Sub-Critical Flow Equation applies (Equation 3.4)
Sub-Critical Vapor Flow
A = [ W /(735 F2 Kd )] [(Z T) / (MW P1 (P1 - P2))]0.5
0.5
A = 1.316 [ W /(735 F2 Kd )] [(Z T) / (MW P1 (P1 - P2))]
0.5
A = 1.342 [ W /(735 F2 Kd )] [(Z T) / (MW P1 (P1 - P2))]
•
3.4 (English)
3.4 (Metric)
3.4 (Metric)
Steam Flow
A = W s / (51.5 P1 Kd Kn Ksh )
A = 1.316 (W s) / (51.5 P1 Kd Kn Ksh )
A = 1.342 (W s) / (51.5 P1 Kd Kn Ksh )
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3.1 (English)
3.1 (Metric)
3.1 (Metric)
3.6 (English)
3.6 (Metric)
3.6 (Metric)
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•
Liquid Trim Valves
0.5
A = [Q / (38 Kd Kw Kv )][G / (P1 - P2 )]
0.5
A = 7.456 [Q / (38 Kd Kw Kv )][G / (P1 - P2 )]
0.5
A = 7.528 [Q / (38 Kd Kw Kv )][G / (P1 - P2 )]
•
3.7 (English)
3.7 (Metric)
3.7 (Metric)
Conventional PSV in Liquid Service
A = [Q / (38 Kd Kw Kv Kp )][G /(1.25Ps - Pb )]0.5
0.5
A = 7.456 [Q / (38 Kd Kw Kv Kp )][G /(1.25Ps - Pb )]
0.5
A = 7.528 [Q / (38 Kd Kw Kv Kp )][G /(1.25Ps - Pb )]
3.8 (English)
3.8 (Metric)
3.8 (Metric)
Note: See Appendix C for a listing of vendor computer programs for valve
sizing
1.4.6
Review Disposal Options (Section 6)
Perform a HAZOP and risk assessment to establish the best location for disposal of
refinery waste streams. See Table 6.1 for typical toxic or hazardous chemicals
encountered around the refinery. The following locations should be considered:
•
Atmosphere
•
Grade or sewer
•
Process Vessel
•
Closed System (Flare Header)
1.4.7
Establish Equipment Depressuring Requirements (Section 4)
1.4.8
Size Thermal Relief Valves (Section 4)
1.4.9
Evaluate Process Flow Loops (Section 4)
•
Settling Out Pressure
•
PSV Relieving rate, if a PSV is required
1.4.10 Evaluate Total Relief Loads to the Flare, by Contingency, to Include Depressuring
(Section 5)
1.4.11 Consider Mitigation for Relief Load Reduction (Section 5)
1.4.12 Review Depressuring Loads for Time Smoothing (Section 5)
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1.4.13 Size Relief Valve Piping Inlet/Outlet (Section 7)
The Relief system piping is to be designed according to the following guidelines:
•
Relief device inlet/outlet piping must meet the pressure drop limitations in
Table 7.1
•
Use the SIMSCI INPLANT program for hydraulics
1.4.14 Establish Required Purging Rates by Converting Velocities Given Below into Rates
3
[lb/h (kg/h) or SCFH (Nm /h)] for the Relief Piping (Section 7)
Note: The quantity of purge gas for an elevated flare depends upon the type of seal
selected.
•
•
Normal Purge Gas Rate:
a.
Elevated Flare - 0.10 ft/s (0.03 m/s) in stack/tip [design for 0.23 ft/s
(0.07 m/s) maximum]
b.
Enclosed Ground Flare - 0.10 ft/s (0.03 m/s) in each first stage tip
[design for 0.23 ft/s (0.07 m/s) maximum]
Special Purge Gas Rate (intermittent):
a.
Elevated Flare - 3.3 ft/s (1.0 m/s) in stack/tip
b.
Ground Flare - 3.3 ft/s (1.0 m/s) in main header
1.4.15 Select and Specify the Following Equipment where Appropriate (Section 8)
o
Blowdown Drum
Required for sizable liquid relief loads only
o
Knock Out (KO) Drum
o
Seal Drum
This drum is mandatory and will be included in every design.
o
Quench Drum
This type of drum is seldom required and is a specialty design item.
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1.4.16 Develop Flare Stack and Tip Details (Section 9):
o
Select flare location
o
Select flare type (reference Table 9.1)
•
Elevated Flare
•
Ground Flare
o
Check if assist fuel is required
o
Select gas seal type
o
Select metallurgy for piping, flare tip and seal
o
Select the flare diameter based on required design flow rate and velocity
limitations and verify the tip diameter proposed by vendor.
o
Calculate the flare stack height based on allowable ground level radiant heat
intensity, using Brzustowski’s and Summer’s method per API RP 521 (see
Section 9.4.4):
•
500 Btu/ft2/h (5,677 kJ/m2/h, 1,356 kcal/m2/hr) for operating areas
•
2
2
2
1,500 Btu/ft /h (17,040 kJ/m /h, 4,069 kcal/m /h) for short term
exposure
•
3,000 Btu/ft /h (34,070 kJ/m /h, 8,137 kcal/m /h) for limited access
areas with shelter
2
2
o
Confirm that acceptable emission limits of Section 9.8.3.3 are met by
selected stack height or adjust height.
o
Check adjacent equipment surface temperatures during design release
(reference Section 9.4.6).
o
Set smokeless flaring rate (reference Section 9.6 and Figure 9.6)
o
Review noise specifications SP-45820 (Equipment Noise Level Limits) and
SP-45230 (Noise Abatement)
o
Review Environmental Emissions to meet local environmental regulations.
o
Establish type of Flare Ignitor:
a.
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Pressure Ignitor
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This type ignitor is traditional for flare systems
b.
Electronic Ignitor
This type ignitor is cheaper and has been providing higher reliability
than in the past. Review vendor designs with special concern for
experience with reliability. This will probably become the standard in
the future.
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Sheet 1 of 6
TABLE 1.1
PRESSURE RELIEF SYSTEM DESIGN RESPONSIBILITIES
Task
Primary
Responsibility
Secondary/Consultant
Responsibility
Notes
Lead Process
Engineer
Unit Process Engineer
Flare System Process
Engineer
Review Scope, Standards and
Guides, and applicable Laws
and Codes for application to the
design. Add client involvement.
System assessment consists of
a review of the process design
to identify where relief system
requirements may affect design
pressures, plot plans or other
basic design specifications. Key
information required: Process
H & M Balances and PFD.
A.) Design
Pressures and
Temperature
Unit Process
Engineer
Control Systems
Engineer
Mechanical Engineer
Piping Engineer
Consider pump/compressor
shutoff heads, additional
throughput, and potential relief
cases.
B.) Relief Stream
Disposal
Unit Process
Engineer
Environmental
Engineer
Define needs for recovery
systems, flares, etc. Identify
any special permit
requirements.
C.) Fire Exposure
Flare System
Process Engineer
Unit Process Engineer
Mechanical Engineer
Review plot plan for location of
vessels and exchangers
containing volatile liquids.
D.) Utility Failure
Basis
Flare System
Process Engineer
Lead Process
Engineer
Establish relief system design
philosophy and utility failure
basis.
E.) System Review
Lead Process
Engineer
Unit Process Engineer
Control Systems
Engineer
Flare System Process
Engineer
Refinery Operations
Mechanical Engineer
Piping Engineer
Review design pressures, utility
failure basis, plot plans and
preliminary relief system
requirements. Identify any
design pressure or layout
changes required or any relief
cases, which need special
attention. Identify any changes
in the utility system plans, which
are beneficial to relief system
design.
1. Establish Scope,
Design
Philosophy and
Standards
2. System
Assessment
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Sheet 2 of 6
TABLE 1.1 (Continued)
PRESSURE RELIEF SYSTEM DESIGN RESPONSIBILITIES
Task
Primary
Responsibility
Secondary/Consultant
Responsibility
Notes
A.) Basic Design
Unit Process
Engineer
Mechanical Engineer
Flare System Process
Engineer
Control Systems
Engineer
Identify sources of process
stream relief loads during
P&ID development, identify
services which require relief
devices.
B.) Design Review
Unit Process
Engineer
Lead Process
Engineer Control
Systems
Engineer
Flare System Process
Engineer
Mechanical Engineer
Perform at time of review of
Issue for Approval P&ID’s.
Review P&ID’s to verify that all
services requiring
overpressure protection have
relief devices and that they are
properly located. Review relief
cases, which should be
evaluated.
C.) Design
Monitoring
Unit Process
Engineer
Lead Process
Engineer Control
Systems
Engineer
Mechanical Engineer
Continue to assure that earlier
design premises remain either
unchanged or within
acceptable limits.
D.) Flare System
Design Activity
Plan
Flare System
Engineer
Lead Process
Engineer
Piping Engineer
Establish flare design activity
plan and identify limitations for
flare. Develop load summary
format and have Unit Process
Engineer fill flare load
summary. Estimate flare load
in reference to past similar
plant for early start of flare
system design.
3. Relief Source
Identification
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Sheet 3 of 6
TABLE 1.1 (Continued)
PRESSURE RELIEF SYSTEM DESIGN RESPONSIBILITIES
Task
Primary
Responsibility
Secondary/consultant
Responsibility
Notes
A.) Preliminary
Relief Load
Calculation
Unit Process
Engineer
Control Systems
Engineer
This step assists in identifying
early needs for adequately
sized PSVs and vessel nozzles
B.) Preliminary
Valve Sizing
Unit Process
Engineer
Control Systems
Engineer
Process engineer performs
sizing calculations.
C.) Preliminary
Nozzle Size
Unit Process
Engineer
Mechanical Engineer
Piping Engineer
Unit Process Engineer
coordinates with vessel group.
Final nozzle size after relief
device selection.
D.) Issue
Preliminary PSV
Data Sheets
Unit Process
Engineer
Flare System Engineer
Control System
Engineer
Issue PSV data sheets to
control system engineer
Unit Process
Engineer
Control Systems
Engineer
Piping Engineer
Mechanical Engineer
Unit Process Engineer
performs a review of inlet &
outlet piping losses. Markup
process P&ID's to reflect
installation.
4. Preliminary PSV
and Vessel
Nozzle Sizing
5. Relief Device
Installation
Review
Piping Engineer performs force
and moment calculations.
Unit Process Engineer and
Piping Engineer to verify that
calculations have been done
and that the installation meets
code and regulatory
requirements.
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Sheet 4 of 6
TABLE 1.1 (Continued)
PRESSURE RELIEF SYSTEM DESIGN RESPONSIBILITIES
Task
Primary
Responsibility
Secondary/consultant
Responsibility
Notes
A.) Flare Load
Summary
Flare System
Engineer
Unit Process
Engineers
Collect unit flare load summary
with relief device calculation
sheets. Check consistency.
B.) Preliminary Flare
System Design
Flare System
Engineer
Mechanical Engineer
Piping Engineer
Perform preliminary sizing
based on collected relief load,
produce P&ID's, size
preliminary flare equipment
consulting with mechanical
engineer, and discuss flare
sterile area and header routing
with Piping Engineer.
C.) Optimize Flare
System
Flare System
Engineer
Unit Process Engineer
Piping Engineer
Obtain flare piping Iso sketches
from piping and perform
detailed flare header hydraulic
analysis. Optimize flare size by
analyzing relief valve
backpressure. Check any
changes of utility failure basis
per project development.
D.) Finalize Flare
Load and Relief
Valve Sizes
Flare System
Engineer
Unit Process Engineer
Piping Engineer
Mechanical Engineer
Finalize flare load and flare
piping to meet the project
objective (cost and schedule).
Inform Unit Process Engineer
and Piping Engineer about the
final size requirement and
buildup backpressure.
Flare System
Engineer
Piping Engineer
Mechanical Engineer
6. Flare System
Design
E.) Issue P&ID's for
Construction
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Issue P&ID's for construction
after client review.
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Sheet 5 of 6
TABLE 1.1 (Continued)
PRESSURE RELIEF SYSTEM DESIGN RESPONSIBILITIES
Task
Primary
Responsibility
Secondary/consultant
Responsibility
Notes
7. Final Relief Load
Computation and
sizing
Unit Process
Engineer
Control Systems
Engineer
Flare System Process
Engineer
Unit Process Engineer
performs final check of relief
load case evaluations and relief
valve sizing based on finalized
flare system design.
8. Final PSV Data
Sheet Revision
Unit Process
Engineer
Control Systems
Engineer
Issue final revision for Control
Systems Engineer to prepare
instrument data sheets and
purchase specifications.
9. “As Purchased”
Equipment
Performance
Review
Unit Process
Engineer
Control Systems
Engineer
Mechanical Engineer
The Unit System Process
Engineers to review relief load
calculations to assess any
effect that “As Purchased”
pump, compressor or control
valve capacities may have
upon relief loads.
10. Monitor design
changes
Lead Process
Engineer
Unit Process Engineer
Control Systems
Engineer
Review any design changes
occurring during project
development and evaluate for
impacts on the design of the
flare and relief system.
Examples would include "As
Purchased " equipment
checks, electrical system
development, cooling water
system design basis and steam
system design basis.
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Sheet 6 of 6
TABLE 1.1 (Continued)
PRESSURE RELIEF SYSTEM DESIGN RESPONSIBILITIES
Task
11. Engineering
Documentation
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Primary
Responsibility
Secondary/consultant
Responsibility
Notes
Unit Process
Engineer
Project Engineer
Lead Process
Engineer
Unit Process Engineer
Prepare relief device
engineering files. Client
Control Systems Engineer and
Unit Process Engineer review
for completeness. Client
project engineer responsible
for requiring that as-built
updates are properly
completed and placed into the
files.
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Sheet 1 of 16
TABLE 1.2
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
1. GENERAL
1. Client relief standards
Provide Applicable Client Engineering
Standards, Standard Specifications and Design
Guides to design team or contractor.
2. Identify flare type required
See Section 9.2
3. Identify criteria for process relief to
atmosphere
All hydrocarbon relief loads vent to relief header.
See ES1100 for reference. See Section 6.5.2.
4. Consideration of flare gas vapor recovery
For significant flare quantities, vapor recovery
economics should be reviewed. See Section
6.8.3.
2. EQUIPMENT/PIPING CRITERIA
Select between elevated or ground flare. See
Section 9.2.
2A. ELEVATED FLARE
1. Design radiation levels
See Section 9.
Per API RP 521. See Section 9.4.5.
Personnel protection
2
2
500 Btu/ft /h (5,677 kJ/m /h, 1,356 kcal/m /h)
Short Term Exposure
1,500 Btu/ft /h (17,040 kJ/m /h, 4,069 kcal/m /h).
Limited access is required around areas where
this limit can be exceeded, for personnel
protection.
2
2
2
3,000 Btu/ft /h (34,070 kJ/m /h, 8,137 kcal/m /h).
Shelter to be available
Restricted Access
2. Spacing for flare stack
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2
Long Term Exposure
2
2
2
See Section 9.4.4. Note: US insurance industry
standards suggest minimum spacing of at least
330 feet (100 meters) from process units,
hazardous storage, tanks and loading or
unloading facilities. This minimum distance
reduces to 213 feet (65 meters) if flare stack
height is greater than 82 feet (25 meters).
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Sheet 2 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
3. Allowable noise
RECOMMENDED PRACTICE
See Section 9.8
50 ft (15 m) from base of flare
Per client standards. Generally, 90 dBA at
maximum smokeless rate and 125 dBA limit
for design flare relief.
At plant boundary
Per client standards. Generally, 75 dBA at
maximum smokeless relief rates
4. Air Pollutants
Must be below the allowable limits specified
in Client Environmental Emission & Effluent
Limits. See Section 9.8.3.3.
5. Smokeless flaring load
Maximum continuous load which can
exceed 5 minutes during any 2 consecutive
hours. If continuous load is not identifiable,
use the smaller of 10% of peak load or flare
tip maximum smokeless capacity given by
the tip supplier. See Section 9.6.
6. Radiation calculation method
Use Brzustowski’s and Sommer’s method
(Sect. 9.4.4).
7. Fraction of heat radiated
0.15 for maximum flare load. 0.25 for
continuous flare load. See Section 9.4.5.
8. Design wind velocity
Maximum wind velocity for maximum flare
load [if none available use 30 ft/s (8.9 m/s)].
See Section 9.4.4.
Normal wind velocity for continuous flare
load [if none available use 10 ft/s (3 m/s)].
9. Maximum flare tip mach number
Max. 0.5 for initial design. Max. 0.7 for
revamp (establish with flare vendor). See
Section 9.4.3.
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Sheet 3 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
10. Flare tip type
See Section 9.7
11. Identify any stack height limitations
See Section 9.4.4
Ground level heat flux
See Item 2A.1 and Client Standard SP46410
Pollutant Dispersion
Must maintain ground level concentration of
air pollution below allowable limits specified
in client standards or government codes .
Aircraft Clearance
Per local aviation regulations
12. Availability of Utilities
Fuel Gas
To be developed on a project basis
Propane
To be developed on a project basis
Compressed Air
To be developed on a project basis
Steam
To be developed on a project basis
Power
To be developed on a project basis
13. Acceptable Seal Types:
Liquid Seal
Flame Arrestor
Gas Seal
Gas Purge (Nitrogen or fuel gas)
Other
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Use at least two out of the listed items, one
of which must be gas purge. Water seal
drums, if used, must be protected against
freezing. See Section 7.5.
See Section 7.5.2 for normal purge rate
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Sheet 4 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
14. Pilot requirements
Continuous pilot required. See Section 9.9
15. Pilot monitoring
Continuous pilot monitoring required. Direct
sensing should be used; thermocouples for
pressure igniters and self-contained system
for electronic igniters. Backup system can
include remote sensor (IR, visual, other).
16. Identify available plot area locations
(per plot plans).
Per project requirements
2B. GROUND FLARE
1. Enclosed Ground Flare
Establish enclosure dimensions
See Section 9.5
Used primarily if normally visible flame is
not acceptable.
A reputable vendor must size the enclosure
Confirm clearance from top of flare
enclosure to adjacent equipment, buildings
or platforms.
Enclosure is sufficiently deep to contain
flames.
2. Other Types
2C. FLARE KNOCK OUT DRUM
1. Select horizontal or vertical vessel
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See Section 9.5
See Section 8.1
Normally a horizontal vessel is selected.
However, spacing and capacity requirements
may dictate a vertical drum.
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Sheet 5 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
2. Sizing Method
See Section 8.1
3. Steam Heater
Required to prevent freezing
4. Pumpout rate criteria
Largest of:
3
3
• 425 ft /h (12 m /h)
• 1.5 x maximum continuous relief rate
• 0.25 x maximum instantaneous relief rate
See Section 8.5
5. Pump Sparing philosophy/driver
Preferred configurations are either one
motor drive with spare turbine drive or two
motor drives with emergency power backup.
Two motor drives supplied with power from
independent low voltage feeders can be an
acceptable alternative, provided that
analysis confirms simultaneous failure of
the two feeders does not also result in
significant liquid relief load. See Section
8.5.
6. Design droplet size
For initial design: 400 micron
For revamp: 600 micron
7. Design liquid gravity
Default value = 0.7 for drum sizing, 0.7-1.0
for pump selection.
8. Design L/D ratio
2:5 to 3:1, unless at maximum drum
diameter for shipping
9. Indicate if any liquid relief services
should discharge to a separate
knockout drum before entering a
common header with vapor relief
loads.
Normally all relief streams are combined in
a common unit header with liquid separation
in a unit knockout drum.
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Sheet 6 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
10. Liquid Surge Capacity
w/auto transfer pump start
w/manual transfer pump start
15 minutes of max. instantaneous liquid
drop out
30 minutes of max. instantaneous liquid
drop out
Where no clear drop out rate has
been established
Horizontal Drum
25 % of the KO Drum diameter
Vertical Drum
20% of the KO Drum Tangent-to-Tangent
length
2D. UNIT BLOWDOWN DRUMS
1. Unit Blowdown Drum to be
provided
2. Sizing Methods
2E. SEAL DRUM
1. Seal leg on the inlet line
2. Seal Depth
Job Specific
Maximum volume relieved over 30 minute
period. See Section 8.2.
See Section 8.3 and Figure 7.3
Minimum 10 ft (3 m) seal leg. See Section
8.3.2.1.
1 ft (0.3 m) minimum or as required to
maintain specified backpressure. See
Section 8.3.2.1.
3. Water above the minimum seal
Sufficient to allow a 10 ft (3 m) vacuum to
be pulled on the flare header without losing
the water seal. See Section 8.3.2.1.
4. Overflow seal height
Greater of 10 ft (3 m) or 150 % of maximum
back pressure. See Section 8.3.2.1.
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Sheet 7 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
5. Winterization
Steam heated coil or nonreactive antifreeze
sealant in the seal drum.
6. Seal water supply
Confirm level is automatically maintained
7. Transfer pump start
Confirm autostart
8. Spare transfer pump
Confirm a spare is available
9. Anti-slosh baffling
Horizontal Drums
Vertical Drums
2F. FLARE HEADER
Add transverse baffles as required. See
Figure 8.4
See Figure 7.3
See Section 7
1. Minimum slope of header
8 in/330 feet (21 cm/100 meters) (0.21%),
per API RP 521. See Section 7.1.1.
2. Design Pressure
Maximum back pressure plus 10 percent or
2
50 psig (3.4 barg, 3.5 kg/cm G), whichever
is greater. See Section 7.1.3.
3. Design Temperature
Stress is based on maximum temperature
possible including fire relief. Heat losses
from the header may be considered. Wall
thickness is based on maximum
temperature excluding fire relief. Minimum
temperature is based on auto-refrigeration
of relief fluids. See Section 7.1.2.
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Sheet 8 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
4. Purge gas source and preferred
method of purging.
Continuous sweep of gas through the
headers using fuel gas or inert gas. See
Section 7.5.
Normal purge gas rate:
• Elevated flare
- Normal operation: 0.10 ft/s (0.03 m/s)
in stack/tip
- Upset condition: 3.3 ft/s (1.0 m/s) in
stack/tip
• Enclosed ground flare
- Normal operation: 0.10 ft/s (0.03 m/s)
in first stage tips
- Upset condition: 3.3 ft/s (1.0 m/s) in
main header
5. Pressure test method
X-ray welds; can also add air/nitrogen test.
See Section 7.1.3.
6. Sonic velocity at branch outlet
acceptable?
Yes, if not in prolonged service. See
Section 7.2.1.6.
7. Define relief header normal
operating pressure w/o emergency
relief loads.
Typical:
2
• 0.43 psig (0.03 barg, 0.03 kg/cm G)
normal [1.0 ft (0.3 m) seal depth]
2
• 1.4 psig (0.1 bar, 0.1 kg/cm G) with
vapor recovery
• Can be higher with enclosed ground
flare
See Section 7.1.7.
8. Define relief header operating
pressure for major relief load
cases.
Per Project Basis
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Sheet 9 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
2G. RELIEF VALVE CRITERIA
RECOMMENDED PRACTICE
See Section 3
1. Indicate any specific limitations on
the use of bellows or conventional
valves.
Normally bellows will be used for relief to
header unless conventional valves are
clearly acceptable based on back pressure.
2. Are pilot valves acceptable?
Indicate any special restrictions or
instructions relating to the use of
pilot valves
Pilot valves will be considered for: light
gases at high pressure where operating
pressures are close to the set pressure, for
system conditions that may cause
chattering in spring loaded valves, and
where remote sensing is required.
3. Indicate any preference or
requirements for locating PSV
directly on vessel nozzle versus on
the overhead piping.
Either option is acceptable. Choice is
based on inlet loss criteria, access and
cost.
4. Indicate any relief services which
are required to be spared and
related inlet and outlet block valve
arrangements.
No spares are normally provided.
Exceptions are equipment governed by
ASME Boiler Code and equipment in critical
services.
5. PSV Installation Criteria
See project design criteria and API RP 520
Part-2
2H. PSV ISOLATION VALVES
Block valves are not normally installed, with
the following possible exceptions:
o Dual type PSV installed on one vessel
o Spare pressure vessel w/PSV used for
standby
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Sheet 10 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
o PSV in parallel w/auto depressuring valve
o Thermal relief valves on piping
Block valves must be car sealed or locked
in open position as required to protect
operating system.
3. ENGINEERING CALCULATIONS
See Section 4
3A. ADMINISTRATIVE EXEMPTIONS
1. All exchangers isolated for
maintenance only are to be
immediately drained.
Eliminates the fire relief case for
maintenance situations. See Section
4.8.3.5.
2. Lock/car seal procedures are to be
rigidly observed and monitored.
This allows maintenance isolation without
adding blocked discharge case.
3B. INDIVIDUAL RELIEF CASES
1. General Fire
See Sections 4.8 and 4.10
a. Applicable Codes:
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General Refinery, Section VIII
vessels.
(Greater than 15
2
psig (1.03 barg, 1.05 kg/cm G)
API RP 520 and 521 (To be confirmed)
Refinery LPG Storage
API RP 520 and 521 with wetted area per
API 2000 for conservative design.
Refinery Low Pressure Storage,
[Less than 15 psig (1.03 barg,
2
1.05 kg/cm G)]
API 2000
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Sheet 11 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
Non-Refinery Low Pressure
Storage, [Less than 15 psig
2
(1.03 barg, 1.05 kg/cm G)]
NFPA 30
Non-Refinery/Gas Plant LPG
Storage
NFPA 58
b. Any special insurance
requirements?
By Client Insurance Underwriter
c. Insulation/banding criteria to
allow credit for insulation.
Stainless steel wire or banding and
aluminum jacketing. See Section 4.8.3.2
d. Insulation credit (F Factor)
0.075 to 1.0, based on calculations per API
guidelines. A value lower than 0.075
should only be used under special
circumstances, based on detailed
calculations and review. See Equation
4.12
e. Include vertical vessel bottom
head in fire heat flux area
calculation
Unless contradicted by specific guidelines
based on skirt design.
2. Air Cooler Fire
Mitigate by free draining and locating away
from fire zone, rather than by relief valve.
3. Exchanger Tube Rupture
Analysis is not required if the low pressure
side test pressure (adjusted for
temperature) is equal to or greater than the
high design pressure. For a typical
hydrotest pressure of 1.5 times design
pressure, this equates to a low pressure
side design pressure greater than or equal
to 2/3 the design pressure of the high
pressure side.
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Sheet 12 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
Where relief protection is provided on the
high pressure side, relief valve set
pressure may be substituted for design
pressure, if lower.
See Section 4.7 for calculation format
4. Air cooler cooling credit during
power failure
25% of normal duty (20 to 30% per API RP
521), provided louvers do not fail closed
due to power failure.
5. Use of instruments to eliminate
individual relief valve load
See Section 5.3 for flare load minimization.
6. Operator response time/criteria to
prevent relief case.
10 minutes allowed. Time begins from alert
from an alarm independent from any
instrument that could cause upset
conditions. Upset must be readily resolved
by one clear operator action. See Section
4.1.7.
3C. UNIT DESIGN BASIS
1. Heat and material balance
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See Section 4.1.1
All loads are based on process design
material balance rates and duties, unless
client requests any specific margin(s) for
future debottlenecking.
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Sheet 13 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
2. Relief Valve Sizing
a. Size relief valve inlet/outlet for
maximum flow based on these
limits, or maximum valve capacity
per requirements in Table 7.1.
b. Turbine drivers rated at 105% of
normal speed.
c. Fired heaters are rated at 125% of
rated firing.
d. Towers operated up to flood or by
reboiler/condenser limits
e. Exchangers in clean condition for
heat addition and fouled condition
for heat withdrawal.
4. RELIEF SYSTEM LOADS
4A. CALCULATE MAJOR FLARE
SYSTEM LOADS
See Section 5
See Section 5.2
1. Fire Area
60 feet (18 meter) diameter fire circle [2500
2
2
to 5000 ft (232 to 465 m ) per API RP 520
criteria] minimum, unless special
containment provided (dikes, walls, etc.)
Entire area of diked storage system
considered to be a single fire circle.
2. Identify Major Utility Power Failure
Modes
Identify any recommended exclusions such
as total power failure. Prepare utility failure
system basis.
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Sheet 14 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
3. Identify impact of power failure on
other utility systems, particularly
cooling water, instrument air, steam
and fuel gas.
Per Project Basis
4. Identify impact of calculated flare
load.
Prepare load summary, calculate radiation
and hydraulic impacts for major loads. If
loads are excessive, identify design
changes or mitigation steps to reduce
loads.
4B. DESIGN MODIFICATIONS TO
REDUCE LOADS
See Section 5.3.2
1. Electrical system configuration
Modify to reduce simultaneous relief loads.
2. Equipment design pressure
Increase design pressure to reduce tower
loads by inducing reboiler pinch.
3. Column accumulator capacity
Increase accumulator volume to avoid
flooding during upset conditions.
4C. RELIEF LOAD MITIGATION
OPTIONS
1. Non-normal automatic
instrumentation (single loop)
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See Section 5.3.3 through 5.3.6
Auto-start spare pumps and single loop
shutdown systems. Only 3 out of first 6
assumed to work. See Table 5.1.
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Sheet 15 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
RECOMMENDED PRACTICE
2. Redundant dual loop trips
High reliability. Four out of 5 assumed to
function on demand. See Table 5.2. Used
normally to cut off heat supply to columns.
3. Redundant loop trips with triple
modular redundant (TMR)
architecture.
Two out of three voting logic. Can be very
expensive and therefore are used only for
most critical services, typically for fired
heaters. Nine out of ten assumed to
function on demand. See Sections 5.3.4.4
and 5.3.5.3.
4. High integrity protective instrument
systems.
Reliability calculated to be in excess of relief
valve reliability. Similar to item 3 above, but
with independent trip sources. Only to be
used where relief valves can not practically
provide protection. Requires calculations to
verify reliability and Client management’s
approval. See Sections 5.3.4.5 and 5.3.5.3.
5. Documentation, maintenance and
testing.
For load reduction credit to be taken for any
of these systems, reliability must be
maintained and proven through testing.
Documentation and quality controls must be
consistent with normal relief system
practices
6. Dynamic simulation
Detailed dynamic simulation can sometimes
provide a basis for significantly reducing
key column loads.
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Sheet 16 of 16
TABLE 1.2 (Continued)
ESSENTIAL CRITERIA FOR FLARE AND RELIEF SYSTEM
ITEM
5. HYDRAULICS
1. Flare header/subheader
RECOMMENDED PRACTICE
See Section 7
Based on worst case (for hydraulics) total
calculated simultaneous releases. See
Section 7.3.
2. Individual laterals
3. Relief valve inlet
Based on relief valve flow used for relief
outlet.
4. Relief valve outlet
See Table 7.1
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FIGURE 1.1
TYPICAL RELIEF SYSTEM
ENGINEERING SCHEDULE
TASK
INITIAL SCOPE (PROCESS TEAM)
A. FRONT END WORK
Basic Process Design
Basic Flare System Design Basis
Rough Equipment Specifications
Relief System Assessment Preliminary
B. PROJECT DEVELOPMENT
P&ID Development (Process)
P&ID Development (Flare)
Relief System Assessment Definitive
C. REVIEW AND DETAILED ENGINEERING
P&ID Review & Approval for Engineering (Process)
P&ID Review/Approval for Engineering (Flare)
Review Relief System Assessment
PSV Location
PSV Sizing (Preliminary)
D. COMMIT DESIGN & EQUIPMENT
P&ID’s Issued Approved for Construction
PSV Final Data Sheet Issue
Relief Load Summaries (Final)
Equipment Purchased
PSV Installation Reivew
Design Changes
Engineering Documentation
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FIGURE 1.2
TYPICAL RELIEF SYSTEM
ACTIVITY FLOW CHART
(1)
Establish Scope
Design Philosophy
and Standards
(2a to 2e)
Relief System
Assessment
(3a to 3d)
Identify Relief
Sources
(4a)
Preliminary Relief
Load Calculation
Prelimary Estimate
(8)
PSV Data Sheet
Revision
(7)
Final Relief Load
Calculations and
PSV Sizing
(5)
Review Relief
Device Installation
(4b to 4d)
Preliminary PSV
and Vessesl Nozzle
Sizing
(9)
Review "As
Purchased"
Equipment
Performance
(6d to 6e)
Final Flare Load
and System Design.
Issue P&ID's for
Construction
(6c)
Optimize Flare
System
(6a to 6b)
Preliminary Flare
System Design
(10)
Monitor Design
Changes
(11)
Engineering
Documents
Note: Numbers in parentheses at the top of the boxes are the Pressure Relief System Design
Responsibilities from Table 1.1
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DESIGN TEMPERATURE AND
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2.0
DESIGN PRESSURE AND TEMPERATURE SELECTION
2.1
DESIGN PRESSURE SELECTION
In order to make an appropriate selection of design pressure it is necessary to understand
the relationship between design pressure and other parameters, such as set pressure, back
pressure, accumulation, MAWP, blowdown, etc. Simple definitions of these parameters
are presented in Section 2.3 and will be discussed in greater detail in later Sections. If the
design pressure is selected without consideration of the relief device requirements then the
design pressure may need adjustment at a later date with negative impacts on project
schedule, cost, and performance of the unit.
2.1.1
Operating Pressure
2
Operating pressure (psig, barg, kg/cm G) is the expected fluid pressure in the
equipment during normal operation; used as a basis for determining design
pressure.
2.1.2
Maximum Operating Pressure
Maximum operating pressure (psig, barg, kg/cm2 G) is the worst case pressure
expected to occur due to process upsets, start-up and shutdown operating cases,
and shut-in operating pressures of compressors and pumps. This is always less
than the design pressure defined below.
The pressure increases are usually caused by equipment characteristics such as
the rise of the pump discharge head caused by higher than normal upstream
pressures and increased pressure rise across the pump due to low flow
(approaching shutoff head), fouling of catalyst beds and filter media area reductions
associated with end of run conditions prior to regeneration or replacement.
2.1.3
Settling Out Pressure
In a reaction loop, the flow of process fluids through a system is achieved by
creating a pressure differential with a pump and a compressor. Where the system
can be shut in and the process flow stopped, the pressure will decrease in the
upstream volumes and increase in the downstream volumes, if the fluid is
compressible. The final pressure is defined to be the settling out pressure and will
be constant throughout the loop after equalization. It is important to establish the
relief device set pressure sufficiently above settling out pressure to prevent flare
relief due to reductions or stoppage of process flow. (See API RP 520 Part I,
Appendix B).
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Design Pressure
2
Design pressure (psig, barg, kg/cm G) is the most severe condition of coincident
internal or external pressure and temperature (minimum or maximum) to be
contained by the equipment. These values are used as the basis for mechanical
design of the equipment. Since allowable design stresses vary with temperature,
the design pressure is always specified with a coincident design temperature.
Some types of mechanical equipment have a design pressure set by conditions
other than those required to contain the maximum anticipated pressure. Where this
is the case, the equipment manufacturer is to be required to provide the maximum
working pressure, the relief device set pressure for this equipment (if one is needed)
and the basis for this required set pressure. For example, compressor casings are
usually based on standard designs which frequently can withstand pressure in
excess of the maximum level anticipated.
Fired heaters are a special case where two design pressures are specified:
•
Elastic Design Pressure
This is the maximum pressure the furnace coil will experience for short
periods of time. This pressure is usually related to pressure relief valve
settings, pump shut-in conditions, etc.
•
Rupture Design Pressure
This is the maximum long-term pressure in the coil during normal operation.
The rupture design pressure is usually the lower of the two design pressures.
When establishing heater design conditions, the process engineer must
therefore identify both the short-term design pressure as well as the
maximum operating pressure, as both figures may be needed to determine
tube wall thickness. For pressure relief system considerations, the shortterm (elastic) design pressure should be considered as the equivalent of
MAWP, and should be selected based on the discussion in Section 2.17.
2.1.5
Design Pressure Selection
2.1.5.1
Pressure Vessels
Select the highest of the following:
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2
1)
Operating pressure plus 25 psi (1.7 bar, 1.8 kg/cm ).
2)
Operating pressure times 1.1 [the margin may be reduced to 1.05 or
a minimum of 100 psi (6.9 bar, 7.0 kg/cm2)], whichever is greater, in
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vessels or reactors with operating pressures over 1000 psig (69 barg,
2
70 kg/cm G). For liquid-full vessels, to avoid frequent liquid relief due
to system hydraulic variations, use 1.2 times operating pressure. For
LPG services, use shutoff pressure times 1.1 to be conservative.
3)
2
50 psig (3.4 barg, 3.5 kg/cm G) if the vessel relieves to the flare.
4)
30 psig (2.0 barg, 2.1 kg/cm2 G) if the vessel relieves to atmosphere.
5)
15 psig (1.03 barg, 1.05 kg/cm2 G) if the vessel is vented to
atmosphere.
The determination of design pressure should be based on the operating
pressure at the top of the vessel. Towers or vessels with significant
pressure drop (top to bottom) should also be specified with a bottom
design pressure.
Vessels which can be subject to a vacuum condition under normal or upset
conditions will be designed for full vacuum. Full vacuum condition resulting
from steamout should also be considered.
All calculated design pressures should be rounded up to the nearest 1.4
2
psi (0.1 bar, 0.1 kg/cm ).
The reactor loop profile and the settling out pressure shall be considered
when setting the design pressure of a vessel in a recycle loop, per API RP
520 Part I, Appendix B.
The process engineer should be aware that selection of the design
pressure actually defines service requirements for the relief system. Often,
a small increase in the design pressure can reduce the cost and complexity
of the relief system or even eliminate the need for pressure relief for
particular contingencies. If the equipment design pressure is low or the
anticipated relieving rates are high, additional care should be taken in this
selection. It should be kept in mind that higher design pressure selection
may reduce or eliminate frequent venting to the flare system.
The flare seal drum, K.O. drum(s) and the flare piping are usually designed
2
for 50 psig (3.4 barg, 3.5 kg/cm G). However, higher pressures shall be
specified if required by hydraulic evaluation.
2.1.5.2
Heat Exchangers
Exchanger design pressure can be set to minimize the need for relief by
selecting the highest of the following, when appropriate:
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1)
If the exchanger is downstream of a vessel, use the design pressure
of the associated vessel plus high liquid level static head.
2)
If the exchanger is upstream of a vessel, use the design pressure of
the associated vessel plus liquid static head, plus line loss at rated
flow from the exchanger to the vessel. Calculate the static head for
two phase flow as if the line is full of cold liquid, unless it is not
possible for the line to fill with liquid.
3)
Design the low pressure side of the exchanger for 2/3 of high
pressure side design pressure if the low pressure side is liquid full..
4)
Maximum pressure from all upstream pumps or compressors, based
on the criteria described in Sections 4.5 and 4.6.
All exchangers in steam service, or subject to vacuum in upset condition,
are designed for full vacuum.
2.1.5.3
Piping
Select the highest of the following, where appropriate:
2.1.6
1)
Vapor from vessel: Vessel design pressure.
2)
Liquid from vessel: Vessel design pressure plus static head from the
high liquid level to grade or lowest elevation plus pressure loss
through vessel internals.
3)
Centrifugal pump or compressor discharge: Maximum pressure as
discussed in Sections 4.5 and 4.6.
4)
Positive displacement pump or compressor discharge: Based on the
relief valve set pressure which is typically set at the higher of the
following:
a)
120% of the rated discharge pressure.
b)
Rated discharge pressure plus 25 psi (1.8 bar, 1.8 kg/cm2).
Design Vacuum
•
Process Equipment
Frequently, vessels in a refinery are designed for significantly greater than
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15 psig (1.03 barg, 1.05 kg/cm2 G) and are capable of sustaining a full
vacuum without further protection. However, many large low pressure
vessels require stiffening rings or additional wall thickness to accommodate
vacuum conditions.
When it is impractical to design low pressure equipment to accommodate
vacuum conditions, the equipment is protected by drawing in atmospheric
air. Low pressure tanks, designed for API Standard 620 or 650, are
examples of this. To prevent the possibility of an unsafe condition caused by
air intrusion, this type of equipment is not vented to the flare headers. This
vacuum design condition will have no impact on the flare and relief system.
Occasionally, equipment must be relieved to the flare, yet it is impractical to
design the equipment for vacuum conditions. In those cases a reliable
source of oxygen free gas is used in lieu of air as a source of pressurization.
•
Flare Equipment and Relief Header
It is possible to develop a vacuum condition in the relief header due to
cooling and condensation of hot relieving vapors by heat loss from the relief
header. This is usually addressed by providing an emergency purge gas
addition when the header pressure declines too much. In any event, the
relief header, seal drum(s) and k.o. drum should be designed for at least a
half vacuum to take care of this case and potential steam out conditions
during start-up and isolation for maintenance. This should cause no
impacts on the vessel and piping designs because of the 50 psig (3.4 barg,
2
3.5 kg/cm G) design pressure of all the flare equipment and piping. (See
Section 7.1.3).
2.1.7
Maximum Allowable Working Pressure (MAWP)
Design pressure and design temperature are used as the basis for design of ASME
Section VIII pressure vessels. A required thickness for the walls is calculated and
the next commercially available size is selected.
The resulting pressure which will appear on the vessel name plate, is identified as
the Maximum Allowable Working Pressure (MAWP). This may provide a margin
which can be used for increasing set pressures and relief capacity, provided all
other components in the system are also suitable for the higher design pressure.
The MAWP always refers to the top or low pressure end of the vessel. Liquid head
requirements and frictional losses must be accounted for in the mechanical design
of the lower sections of the vessel.
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2.2.1
2.2.2
Definition of Operating Temperature
2.2.1.1
The Normal Operating Temperatures are the fluid temperatures which
appear on the Process Flow Diagram.
2.2.1.2
Abnormal Operating Temperature is the fluid temperature in the equipment
during unusual operations or process upsets. The following conditions are
not normal operating conditions, generally. These conditions should be
reviewed on a case by case basis considering the frequency and duration
in order to avoid excessive investment. If these conditions should be
implemented to design conditions, it is necessary to specify multiple design
temperatures coincident with pressure to establish a proper design.
1)
Operation at start-up, or shut-down when off-spec products are
produced.
2)
Operation bypassing a heat exchanger or a vessel which is normally
operated including loss of coolant.
3)
Catalyst regeneration, heater decoking, steam purge, and nitrogen
purge operation.
4)
Emergency depressuring and relieving of safety relief valve.
Maximum Operating Temperature
Maximum operating temperature is the worst case temperature occurring during
normal operation, process upsets, start-up, and shutdown operations. This is
usually less than the design temperature defined below. An exception would occur
if it is impossible to reach the design pressure while at the maximum operating
temperature and this combination of temperature and pressure do not control any
aspect of the mechanical design.
2.2.3
Definition of Design Temperature
2.2.3.1
Design Temperature
Design temperature is the highest temperature expected to exist at the
same time as the design pressure. It is not unusual to specify the design
temperature as the maximum value that will not cause a decrease in
allowable design stress. This is usually 650 °F (343 °C) for carbon steel.
However, it should be made clear to the detailed design team and to
operations that this temperature was selected as a matter of convenience
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and is not to be used as the basis for pipe stress calculations for the
equipment or connected piping or for increasing operating temperature.
See Figure 2.2 for a graphical representation of this allowable temperature
and design stress relationship.
2.2.3.2
Minimum Design Metal Temperature (MDMT)
The MDMT is required on all pressure vessels which are to carry an ASME
code stamp, including heat exchangers. The same guidelines discussed
below will apply to all code-stamped equipment.
The MDMT is the lowest temperature of the contents of the pressure
vessel during one of the following two conditions:
Condition 1:
Normal operation, startup or shutdown
For this condition, the MDMT should be set
approximately 10 °F (6 °C) below the lowest operating
temperature occurring during normal operation, startup, shutdown, regeneration, etc., considering expected
variations in composition and SOR/EOR conditions.
The MDMT may be set at less than 10 °F (6 °C) below
the operating temperature determined above with the
approval of the lead process engineer. This situation
may occur when there is a substantial economic
advantage and system safety can be maintained.
Condition 2:
Expected off-design conditions
If the vessel can reach equilibrium with the ambient
temperature under normal operation (i.e. pressurized
storage vessels), then the MDMT shall not be higher
than the site minimum design temperature and shall
be set concurrent with the vessel design pressure.
A MDMT corresponding to the lowest one-day mean
ambient temperature coincident with the vessel design
pressure shall be used if this situation provides the
lowest MDMT.
If depressuring the normal contents of a vessel can
result in auto-refrigeration, this phenomenon must be
addressed. The lead process engineer should work
with the mechanical and metallurgical engineers to
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determine the design-controlling case of coincident
temperature and pressure.
The lead process engineer must approve any
deviations from the above guidelines.
2.2.4
Design Temperature Selection
Add 50 °F (28 °C) to the maximum operating temperature (paragraph 2.2.2)
coincident with the design pressure case (paragraph 2.1.3 and 2.1.4) for
temperatures up to 650 °F (343 °C) and plus 25°F (14 °C) over 650 °F (343 °C) to
obtain the design temperature. Remember, this maximum operating temperature is
to be simultaneous with the design pressure. This allowance is to provide for
variations from normal operating conditions caused by temperature controllers and
process fluctuations. If greater deviations can be foreseen, the larger anticipated
value should be used. Such deviations can occur for example, as a result of upset
conditions. Bubble point temperatures in towers at relief conditions and reactor loop
exchanger temperatures during depressuring are examples of conditions which
should be considered.
Design temperatures for columns shall be based on maximum internal operating
temperature. However, two or more design temperature zones should be specified
if economic considerations dictate.
2.3
PSV RELATED PRESSURES
Proper selection of design pressures and specification of relief valves requires a clear
understanding of the terms in this section. The concept of setting design pressure at 10%
over normal operation is correct most of the time but an understanding of when exceptions
can be made and the basis for these exceptions will become clear from an understanding of
these concepts.
See Figure 2.1 for a graphical interpretation of these pressures.
2.3.1
PSV Set Pressure for Vessels
Where the MAWP is greater than the design pressure, and the entire system design
is suitable for the higher pressure and possible higher temperature, it is a good
practice to set the relief valve at the MAWP. This will allow additional operating
flexibility without increased equipment cost. Some advantages would be increased
allowable tolerance related to set pressure, reduced orifice size requirement for a
given flow and reduced potential for relief.
If the MAWP is unknown at the time the PSV is being specified, the set pressure
shall not exceed the equipment design pressure.
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For multiple relief valves the set pressure of one valve must be set at or below
equipment MAWP. The following criteria apply to the establishment of set pressure
for the remaining valves:
1.
Where an additional pressure relief valve is provided for fire, only that relief
valve can be set at 10% above the MAWP.
2.
For ASME, Section VIII vessels or boilers, set the one valve at MAWP with
remaining valve(s) set at up to 105% of MAWP.
3.
For boilers designed to meet ASME Code Section I, the highest set pressure
is 3% over the design pressure or MAWP.
System hydraulic losses must be considered at incipient relief and fully accumulated
relieving conditions to assure that all prevailing pressures are within code-allowable
limits.
2.3.2
Spring Setting (Cold Differential Test Pressure)
The spring settings of spring loaded pressure relief valves are not always equal to
the set pressure. Adjustments are sometimes needed for the superimposed back
pressure and temperature.
2.3.3
Permissible Overpressure or Accumulation
When a conventional relief valve opens against its spring setting, the open area for
relief increases as the pressure rises over that required to first unseat the valve.
The valve will continue to open until the required relief rate passes through the
valve. The increase in pressure over the valve set pressure is overpressure.
Where the set pressure equals the MAWP, this is the same as accumulation.
Accumulation is the same absolute pressure rise but is calculated as a percentage
of vessel MAWP.
The amount of permissible accumulation depends on the design code used for the
protected equipment:
•
ASME Section VIII Pressure Vessel Code
°
°
°
•
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MAWP plus 10%
MAWP plus 16%
MAWP plus 21%
- Single valve (non-fire case)
- Multiple valves (non-fire case)
- Fire case
ASME Section I Boiler Code - MAWP plus 6%
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Superimposed Back Pressure
The static pressure at the discharge of the relief device at the time the device is
required to operate is the superimposed back pressure and may be very close to
atmospheric pressure or significantly higher. This pressure could have a large
impact on the ability of the relief device to open when set pressure is reached and in
extreme cases could reduce the capacity of the valve when it does open. It is very
important to understand the effects of high pressure at the relief device outlet on its
operation in order to assist in selection of the correct relief device. This
superimposed pressure may vary due to changes in operating conditions on the
discharge side.
2.3.5
Built-Up Back Pressure
This is the increase in pressure at the relief device outlet that develops as a result of
flow after the pressure relief device opens. Note that it has units of differential
2
pressure, psi (bar, kg/cm ).
2.3.6
Back Pressure
Back pressure is the sum of the superimposed and built-up back pressure,
psig (barg, kg/cm2 G).
Note that API RP 520 requires that the combined effect of the superimposed and
built-up back pressure on the performance characteristics of the valves must be
considered when more than one pressure relief valve discharges into a common
manifold simultaneously.
2.3.7
Pressure Tolerances
The tolerance for set pressure is +/- 3% according to API RP 521. This means that
the relief device will open within 3% of the set pressure. As an example, for a 142
psig (9.8 barg, 10 kg/cm2 G) set pressure, the device is considered to be fully
functional if it begins to open between 138 (9.5, 9.7) and 146 psig (10.1 barg, 10.3
2
kg/cm G).
The tolerance for blowdown and simmer (initial flow of vapor before full lift) are not
specified by ASME Section VIII.
Tightness tolerance is defined in API RP 527. For conventional relief valves, a
normal tolerance for tightness is 10 % of set pressure but increased tightness can
be purchased up to about 95% of set pressure. Pilot operated relief valves can be
expected to remain tight at operating pressures up to 95% of set pressure.
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Blowdown Pressure
This is the differential pressure required to reseat the relief valve after opening. It is
usually about 2% to 4% for boilers (by code) and 5% to 10% for ASME Section VIII
pressure vessels.
2.4
EQUIPMENT RERATING
When changes in process conditions are required, such as increased operating pressure or
temperature, it may be possible to rerate the equipment rather than provide a new or
modified relief system. Rerating must be carefully considered because of the design
uncertainties, potential for increased project duration and increased costs that can result.
Good inspection records of wall thickness will help minimize the uncertainties and reduce
the time to implement a rerating. ASME vessels have to be hydrotested at up to 150
percent of the new MAWP and the code papers and refinery inspection files will need to be
updated. Restamp any upgraded vessels with the new MAWP. Mechanical equipment
rerating requires calculations from the equipment manufacturer or another company
certified to rerate and restamp under the applicable code.
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FIGURE 2.1
TYPICAL PRESSURE LEVELS PER API RP 521
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FIGURE 2.2
ALLOWABLE DESIGN STRESS VERSUS TEMPERATURE
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3.1
INTRODUCTION
RELIEVING DEVICES
Proper selection and specification of pressure relief devices is as important to pressure relief
system design as is proper evaluation of potential relieving cases and load calculations. Most
pressure relief devices are simple mechanical devices which are operated by the energy
contained in the relieving fluid. Their designs favor simplicity and reliability over flexibility.
Because they are generally operated only by process energy, most pressure relief devices
have operating characteristics and limitations which are not found in other types of automatic
valves and which need to be recognized during design.
3.2
TYPES OF PRESSURE RELIEF DEVICES
There are two broad classes of pressure relief devices. The first class relates to those devices
which are constructed and tested in accordance with the various sections of the ASME Boiler
and Pressure Vessel Code. These devices are used in services with set pressures exceeding
15 psig (1.03 barg, 1.05 kg/cm2 G). The second class involves those devices which are not
covered by the ASME Code. Generally, but not always, these are devices with set pressures of
2
15 psig (1.03 barg, 1.05 kg/cm G) or less.
Within the ASME device class, there are reclosing and non-reclosing pressure relief devices.
In refinery services, the reclosing devices are all pressure relief valves. In refinery services,
non-reclosing types are effectively only rupture disks. Other types of non-reclosing devices
exist, however their use in refinery services is extremely rare.
The non-ASME class is primarily those devices with set pressures of 15 psig (1.03 barg, 1.05
2
kg/cm G) or less. In refinery services, these generally are low pressure storage tank vents and
breathers and steam turbine surface condenser pressure relief valves. Some non-ASME
devices which may have set pressures above 15 psig (1.03 barg, 1.05 kg/cm2 G) include
sentinel valves mounted on steam turbine cases and pressure relief valves installed on services
outside the scope of the ASME code.
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Following are general descriptions of the various types of pressure relief devices commonly
used in refinery service. Detailed descriptions of their operational and application limitations
follow, beginning in Section 3.4.1.
3.2.1
Safety Valves
Safety valves are spring opposed pressure relief valves characterized by a rapid lift
from closed to a nearly wide open position when the set pressure of the valve is
reached. This is termed a “popping “ action. Generally, safety valves are used only in
steam generator services covered by Section I of the ASME Boiler and Pressure Vessel
Code. Safety valves are designed for use only on vapor services and may not be used
in liquid service.
3.2.2
Relief Valves
Relief valves are spring opposed pressure relief valves characterized by a lift that is
proportional to the overpressure above set pressure. Relief valves do not exhibit a
popping action such as a safety valve does. Full lift and capacity generally are not
reached until the inlet pressure reaches 125% of set pressure. Prior to 1985, relief
valves were often used in liquid services. Since 1985, use of relief valves has been
severely restricted and they may not be used in new construction services which are
covered by the ASME Boiler and Pressure Vessel Code or in services which require
ASME code stamped pressure relief valves.
3.2.3
Safety Relief Valves
Safety relief valves are spring opposed pressure relief valves which have the operating
characteristics of a safety valve when operating on a compressible fluid (vapor or a
multi-phase vapor/liquid stream) and the operating characteristics of a relief valve when
operating on a liquid stream. Until 1985, safety relief valves were universally used for
virtually all services covered by Section VIII of the ASME Boiler and Pressure Vessel
Code. Applications after 1985 are restricted to vapor and vapor/liquid services. Within
this restriction, and except for steam generator safety valves, virtually all pressure relief
valves found in refinery services are safety relief valves.
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RELIEVING DEVICES
Liquid Trim Relief Valves
This section only applies for services where liquid relief is controlling. For services
where vapor relief is controlling, use vapor relief valves.
The 1986 edition of Section VIII of the ASME Code required that all pressure relief
valves in liquid service be designed and certified to achieve full lift at an inlet pressure of
no more than 110% of set pressure. As existing safety relief valves could not meet this
requirement, a new design, the liquid trim relief valve, was required. For applications
installed after January 1, 1995, liquid trim relief valves meeting these requirements must
be used in all liquid services which fall under the scope of the ASME Code.
Some manufacturers have successfully certified their liquid trim relief valves for
operation in compressible fluid services. However, vapor operating characteristics,
particularly blowdown, may be different than for safety relief valves, and not all liquid
trim relief valves may be used in compressible fluid services.
3.2.5
Pilot Operated Pressure Relief Valves
Pilot operated pressure relief valves are a type of pressure relief valves which is
activated by a small auxiliary spring opposed pressure relief valve, called a pilot. When
the system pressure is below set pressure, the pilot routes the process pressure to a
chamber, called the dome, above a piston or diaphragm which holds the main valve
closed. When the set pressure is exceeded, the pilot vents the process pressure from
the dome and allows the process pressure at the inlet of the valve to push it open.
Pilot operated pressure relief valves may be used in either vapor or liquid service and
are typically used when the normal operating pressure exceeds 90% of set pressure, or
when the location at which pressure must be sensed is different than the location at
which relief is to occur.
Pilot operated pressure relief valves are available in many configurations, and certain
components have operating limitations. Engineering judgment and a clear
understanding of system requirements are needed in applying these type of valves.
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Rupture Disks
Rupture disks are non-reclosing differential pressure relief devices which are designed
to burst or tear open when the differential pressure across the disk exceeds a specific
value. A rupture disk device includes a rupture disk and a rupture disk holder. (Section
3.9 describes the various types of rupture disks and their applications.)
3.2.7
Non-ASME Devices
Pressure relief devices which are not constructed and tested in accordance with ASME
Code are termed “Non-ASME” devices, and may be used only in those services which
are not governed by the ASME Code. In practice, these situations are encountered in
two areas:
2
o
Services with set pressures of 15 psig (1.03 barg, 1.05 kg/cm G) or less.
o
Mechanical equipment or other services which are outside of the scope of the
ASME Code, even if the device set pressure is above 15 psig (1.03 barg, 1.05
kg/cm2 G).
Many of these types of devices are designed for special purpose applications. Devices
for steam turbine applications are described in Section 9.11. Devices for use in
protecting low pressure storage tanks are described in Section 8.0.
Other types of devices do exist, but are seldom encountered in refinery applications.
When such a requirement is identified, the engineer should consult with the
manufacturer of the device and the manufacturer of the protected equipment to ensure
that the relief device is being properly specified and installed and that all regulatory
requirements are being met.
3.3
CODES AND STANDARDS
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With the exception of low pressure [less than 15 psig (1.03 barg, 1.05 kg/cm2 G)] relief devices,
the design and certification of virtually all pressure relief devices is governed by ASME Code.
The ASME Code specifies basic requirements for construction, set pressure, basic device
performance, and capacity testing and certification. Other standards also exist which govern
pressure relief device application, but these documents usually are based upon use of ASME
approved devices.
3.3.1
ASME Section I
ASME Section I provides minimum specification, construction, testing and certification
requirements for fired boilers and organic vaporizers. Section I only addresses safety
valves and safety relief valves since other types of relief devices are not used for
Section I application.
3.3.2
ASME Section VIII
ASME Section VIII provides minimum specification, construction, testing and
certification requirements for pressure relief devices installed to protect ASME Section
VIII vessels from overpressure. Applicable Code sections for Section VIII pressure
relief devices include paragraphs UG-126, UG-127, UG-129, UG-131, UG-132 and UG136.
ASME Code requirements for rupture disk burst pressures are the same as for pressure
relief valve set pressures. Code requires that the stamped pressure of a rupture disk
used in combination with a pressure relief valve not be higher than the relief valve set
pressure.
3.3.3
ANSI/API Standard 526
ANSI/API Standard 526 defines basic construction, external dimensions and pressure
and temperature limits for flanged steel safety relief valves. This standard does not
provide certification or functional requirements as Section VIII of the ASME Code does,
but serves to standardize designs of a class of ASME approved safety relief valves.
Standard 526 addresses the following items:
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o
RELIEVING DEVICES
Standard pressure relief valve effective orifice areas. These orifices are
identified by the letter designations commonly used to specify pressure relief
valve nozzle sizes.
o
Standard face to face dimensions for pressure relief valve bodies.
o
Definition of standard combinations of nozzle sizes and inlet and outlet flange
sizes.
o
Definition of standard materials of construction.
o
Defi nition of standard limitations on set pressures, back pressures and service
temperatures for each inlet/outlet/orifice area and material combination.
ANSI/API Standard 526 is not a mandatory standard, but is adopted by many users as a
means of achieving common pressure relief valve capacities and dimensions. Virtually
all flanged spring opposed safety relief valves are constructed in accordance with API
Standard 526. However, API Standard 526 does not address threaded body pressure
relief valves or safety valves.
3.3.4
API RP 520, Part I
API Recommended Practice 520, Part I addresses sizing and selection of relief devices
and Part II describes recommended installation practices.
3.3.5
Testing and Certification
ASME Code provides specific methods for testing and certification of pressure relief
devices by the manufacturers. In addition to basic Code requirements, ASME
PTC/ANSI Standard 25.3-1976 defines test equipment and methods for testing
pressure relief valves.
Section VIII of the ASME Code provides methods for either computing the capacity, or
testing the capacity of rupture disks (paragraph UG-127 (a)(2)). If a rupture disk is not
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capacity tested, ASME Code requires that the maximum discharge coefficient which
may be used in determining the rupture disk capacity is 0.62.
3.3.6
Code Stamps
For a safety, safety relief, pilot operated or liquid trim relief pressure valve to bear an
ASME Code stamp, it must be designed, manufactured, tested and certified in
accordance with the requirements of the ASME Code. ASME Code provides definition
of what information must be included on the valve nameplate (UG-129 for Section VIII,
and PG-110 for Section I).
It is important to note that the manufacturer is not required to provide ASME or NB
stamps on ASME approved valves unless it is required in the purchase specification.
There is no provision in the ASME code for rupture disks to bear an ASME code stamp.
UG-129(e) defines the marking requirements for rupture disks when in service alone,
and UG-129(c) defines marking requirements when a rupture disk is used in
combination with a safety or safety relief valve.
3.4
CONVENTIONAL PRESSURE RELIEF VALVES
This section describes how conventional safety and safety relief valves operate and discusses
their characteristics and application limitations. Balanced bellows safety relief valves are
described in Section 3.5 and liquid trim relief valves are discussed in Section 3.6.
Figure 3.1 shows a cross section of a typical conventional pressure relief valve. The valve
consists of an inlet nozzle, a disk which is held against the nozzle to prevent flow under normal
system operating conditions, and a spring which provides the force to hold the disk against the
nozzle. The valve set pressure is determined by adjusting the compression of the spring.
When the system pressure at the inlet of the valve overcomes the downward force of the
spring, the valve will open.
Trim designs vary, depending upon the Code to which the valve is constructed. Section VIII
safety relief valves are required to be fully open when the inlet pressure is 10% above the valve
set pressure, and do not have specific reclosing pressure requirements (blowdown). Section I
safety valves have much more stringent requirements, with the valve having to be fully open
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when the inlet pressure is 3% above set pressure, and to reclose at no less than 4% below set
pressure.
3.4.1
Operating Characteristics
The operating characteristics of a conventional safety relief valve are shown
diagrammatically in Figure 3.2-A. At a pressure below the set pressure (typically 93 to
98% of set pressure, depending upon valve maintenance and condition), some slight
leakage (“simmer”) may occur between the valve seat and disk. This is due to the
progressively decreasing net closing force acting on the disc (spring pressure minus
internal pressure). As the operating pressure rises, the resulting force on the valve disk
increases, opposing the spring force, until at the set pressure (normally adjusted to
equal the vessel design pressure) the forces on the disk are balanced and the disk
starts to lift.
As the vessel pressure continues to rise above set pressure, the spring is further
compressed until the disk is at full lift. The valve is designed to pass its rated capacity
at the maximum allowable accumulation.
Following a reduction of vessel pressure, the disk returns under the action of the spring
but reseats at a pressure lower than set pressure by an amount termed the blowdown
(4 to 8% of set pressure). The relief flow just prior to reclosure is typically around 25
percent of valve rated capacity. The blowdown may be adjusted within certain limits, by
various means recommended by the valve vendor or manufacturer, to provide a longer
or shorter blowdown.
Pressure relief valves for vapor service (i.e., safety valves and safety relief valves) are
specifically design for “pop” action (see Figure 3.2-A). That is, they move to the full
open position at only a slight overpressure, the valve remaining full open as
overpressure builds up to the permissible maximum, at which condition the rated
quantity is discharged. Typically, this initial opening may result in 50-60% of rated
capacity flow. This “pop” characteristic is achieved by a secondary annular orifice
formed outside the disk-to-nozzle seat. This causes additional disk area to be exposed
to the operating pressure as soon as a slight lift occurs, accelerating the opening
movement.
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The kinetic energy of the flowing vapor, by action between the valve disk holder and the
blowdown ring, adds to the opening force and causes the valve to “pop” open. This
flowing kinetic energy also continues to act against the spring force as the fluid pressure
returns to the pressure relief valve setting. This accounts for the fact that the pressure
relief valve reseats at a lower pressure than the set pressure; i.e., blowdown.
As normally designed, vapor flow through a typical high-lift safety relief valve is
characterized by limiting sonic velocity and critical flow pressure conditions at the orifice
(nozzle throat), and for a given orifice size and gas composition, mass flow is directly
proportional to the absolute upstream pressure.
Pressure relief valves in liquid service (i.e., relief valves and safety relief valves) have
the characteristic of progressively increasing lift with rising inlet pressure until the full
open position is reached at about 25% overpressure (see Figure 3.2-B). This
characteristic may vary between types and between make. Pressure relief valves in
liquid service may be unstable at flows below about 40 percent of rated capacity,
causing the valve to chatter.
3.4.2
Applications
Safety valves are used in steam services which are under the jurisdiction of Section I of
the ASME Code. Safety relief valves are used for virtually all other vapor or two-phase
(vapor/ liquid) services. Section VIII of the ASME Code defines the basic functional and
construction requirements for safety relief valves used in ASME services.
Conventional spring opposed pressure relief valves are used in virtually all relief
services which discharge to atmosphere or to a constant pressure system. They should
not be used in applications which have variable back pressures or which could have
built-up back pressures in excess of the valve’s tolerance.
3.4.3
Design Considerations
Spring opposed pressure relief valves are highly dependent upon the force balance
between the valve spring and forces generated by the flowing fluid. This section
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describes the important design considerations that must be observed in specifying
spring opposed pressure relief valves. These include:
o
Operating pressure generally ≤ 90% of set pressure.
o
Superimposed backpressure, if any, is constant
o
Inlet frictional losses ≤ 3% of set pressure at valve rated flow.
o
Built-up backpressure ≤ 10% of set pressure for safety relief valves at valve
.
rated flow (at 10% accumulation).
3.4.3.1 Operating Pressure
Spring opposed pressure relief valves are prone to leakage if the normal
operating pressure is too close to the set pressure. For standard design spring
opposed pressure relief valves, the maximum normal operating pressure should
not exceed 90% of set pressure. If this value is exceeded, the valve could leak
a small amount of fluid continuously. In some cases, particularly with steam, a
small leak will cause wire-drawing of the valve seating surfaces, causing the leak
to get even larger.
A second problem with operating pressure relief valves close to their set
pressures is that minor process upsets may result in the set pressure of the
valve being exceeded, causing the valve to pop. In addition to the nuisance
factor in having pressure relief valves popping due to small process upsets, the
valve may not properly reseat, which will require it to be removed from service to
be repaired.
It is possible to specify spring opposed pressure relief valves for normal
operating pressures greater than 90% of set pressure. When this is done, the
manufacturer may require special honing of the nozzle and disk seating surfaces
to achieve an acceptable leakage rate or may equip the valve with a soft seat.
However elastomers used as soft seating materials may not be suitable for the
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operating temperature or process fluid. Either of these approaches requires
special maintenance procedures and should not ordinarily be utilized.
3.4.3.2
Superimposed Back Pressure
Conventional pressure relief valves usually are used when the discharge of the
valve is routed to atmosphere through a short tail-pipe. It is possible to use
conventional valves in installations where the back pressure is other than
atmospheric, but the effect of the destination system back pressure must be
compensated for in setting the valve set pressure.
3.4.3.3
Inlet Loss
The operation of a spring opposed pressure relief valve is dependent upon the
energy of the fluid flowing though the valve providing the force necessary to hold
the valve open. If the piping frictional losses between the vessel being protected
and the pressure relief valve are too high, there may not be sufficient energy in
the fluid to keep the valve open.
A pressure relief valve starts to open at its set pressure, but under discharging
conditions, the pressure acting on the valve disks is reduced by an amount
equal to the pressure drop through the inlet piping and fittings. If this pressure
drop is sufficiently large, the valve inlet pressure may fall below reseating
pressure, causing it to close, only to reopen immediately since the static
pressure is still above the set pressure. Chattering results from the rapid
repetition of this cycle.
To avoid this mechanism as the cause of chattering, inlet piping to pressure
relief valves should be designed for the lowest practical inlet pressure drop
(including exit loss and piping and isolation valve pressure drop). API RP 520
guidelines as well as manufacturers’ recommendations dictate an inlet pressure
drop of no more than 3% of set pressure at maximum valve relieving capacity.
This 3% limit is particularly important for pressure relief valves in liquid service.
In the case of pressure relief valves in low pressure vapor services where the set
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pressure is below 50 psig (3.4 barg, 3.5 kg/cm2 G), or for certain valve sizes
(particularly larger valves), this may be difficult to achieve and the manufacturer
should be contacted for guidance. Loss up to 5% may be acceptable for certain
valves.
3.4.3.4
Back Pressure
The back pressure imposed on a pressure relief valve consists of superimposed
back pressure present prior to relief and built-up back pressure which occurs as
a result of fluid flowing through the valve and its outlet piping.
While the effects of constant superimposed back pressure can be compensated
for by adjusting the cold differential set pressure, the effects of built-up pressure
cannot be. As with constant back pressure, the effect of built-up back pressure
in a conventional spring opposed pressure relief valve is to increase the total
down force on the valve disk. However, this force is only present when the valve
is flowing, and does not interfere with the valve’s set pressure.
If built-up back pressure is high enough, the additional force imposed during
flowing conditions will be sufficient to overcome the upward forces on the disk
and force the valve closed. This will cause valve chatter in a manner similar to
that which occurs if inlet losses are too high. Once the valve has been forced
closed by built-up back pressure, the flow causing the build-up stops, the force
under the valve is sufficient to re-open it, and the cycle starts over again.
Outlet piping for conventional spring opposed safety relief valves should be
sized for a pressure drop not exceeding 10% of set pressure based on the
valve’s rated low at 10% accumulation. (Note: This is true even when valves are
sized for 16% and 21% accumulation).
If total back pressure is greater than the calculated critical flow pressure, the
capacity of a conventional pressure relief valve is affected by the back pressure
since flow will be subsonic. Back pressure reduces the capacity of conventional
pressure relief valves in liquid service, which depend on pressure differential for
flow.
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Safety valves have the area above the valve disk vented to atmosphere, so the
force balance across the disk is not affected by built-up back pressure in the
same manner as for safety relief valves. However, the special disk designs
required for safety valves to meet Section I Code requirements makes their
performance also sensitive to built-up back pressure. Most manufacturers
recommend that their safety valves not be subjected to any more than 10% builtup back pressure. It should be noted that in this case, the 10% back pressure
limit applies when the safety valve is flowing at rated capacity at 3% over
pressure.
3.5
BALANCED BELLOWS PRESSURE RELIEF VALVES
Balanced bellows pressure relief valves have the same basic construction, and operate in the
same manner as conventional spring opposed pressure relief valves. However, these valves
incorporate means for reducing the effect of back pressure on the set pressure and the effect
of built-up back pressure on the valve’s performance characteristics. Balanced bellows
pressure relief valves come in two basic designs; the piston type and the bellows type.
3.5.1
Operating Characteristics
The piston type is seldom specified for use in the refining industry because flare header
gases would continuously vent pass the piston and its guide and the small clearance
between the piston and guide provide a site for material accumulation which can result
in binding.
The bellows type, shown in Figure 3.3, has a flexible bellows installed around the disk
holder and spindle that is vented to atmosphere. The effective bellows area is the same
as the nozzle seat area. The balanced areas prevent the back pressure from acting on
the top side of the disk within the area covered by the bellows. The disk area extending
beyond the effective bellows area and the opposing nozzle seat area are equal which
cancels the effect of the back pressure on the valve disk so that there are no
unbalanced forces. In practice, the cancellation is not perfect and at higher back
pressures the valve capacity is affected. This can be compensated for when the valve
is sized.
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The bellows also serve to isolate the disk guide, spring and other top-work parts from
the relief fluid and any fluid in the discharge piping. This feature can be important in
refinery services where the relief fluid can be corrosive or may foul the pressure relief
valve.
Due to physical size limitations, balanced bellows are not available in certain valve
designs and sizes. In these cases, unbalanced bellows valves may be specified where
only corrosion isolation is required. In some other cases, such as occurs in some
manufacturer’s smaller valves (D and E orifices), a larger orifice with a restricted lift (to
reduce capacity back to that of the smaller orifice) may be used in order to balance the
valve.
3.5.2
Applications
Bellows valves should be specified where any of the following apply:
o
Superimposed back pressures are not constant. (Where back pressures
fluctuate on a conventional valve, the valve may open at too low a pressure or
permit the vessel pressure to exceed the equipment rating, depending upon
back pressure fluctuation).
o
The built-up back pressure exceeds 10% of the set pressure (at 10%
accumulation).
o
The service is fouling or corrosive, since the bellows shields the spring from
process fluid. Note, however, that the bellows convolutions could also foul in
extremely viscous service, such as asphalt, limiting the lift of the valve unless the
valve is heated and insulated.
o
3.5.3
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Preference should be given to balanced bellows valves whenever discharging
into a flare header, closed collection system, process vessel or process piping.
Design Considerations
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Balanced bellows pressure relief valves may be used satisfactorily in vapor and liquid
service with a back pressure (superimposed plus build-up) as high as 30% of set
pressure. Higher back pressures are allowable when corrected for overpressure and
flow capacity derating. The back pressure must be incorporated in the sizing
calculation.
In retrofits it may be acceptable to exceed 50% of set pressure. In such cases the valve
manufacturer should be contacted to establish the reduction in capacity due to the high
back pressure for the particular pressure relief valve involved. In no case should the
back pressure exceed 75% of set pressure.
In addition to the above back pressure limitations based on valve capacity, balanced
bellows pressure relief valves are also subject to back pressure limitations based on the
mechanical strength of the bellows or bellows bonnet, or the valve outlet flange rating.
The back pressure specified for the valve is governed by the lowest back pressure
permitted by these various criteria.
Although the bellows pressure relief valve has the advantage of tolerating a higher back
pressure than the conventional valve, it should be recognized that the bellows is
inherently a point of mechanical weakness which introduces some degree of additional
risk, in case the bellows should fail and release process fluids through the vent. They
should be avoided in services where the process temperature exceeds the auto-ignition
point.
In order to achieve the required balancing of the valve disk, the interior of the bellows
must be vented through the bonnet chamber to the atmosphere. A vent hole is provided
in the bonnet for this purpose. Thus, any bellows failure or leakage will permit process
fluid from the discharge side of the valve to be released through the vent. Venting
arrangements must be carefully considered to avoid entry of rain water, impingement on
lines or equipment, or exposure of personnel to hazardous material.
Non-balanced bellows design safety relief valves have the same back pressure
restrictions as conventional safety relief valves. All bellows or piston type safety relief
valves are also restricted to inlet losses just as conventional safety relief valves are.
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3.6
RELIEVING DEVICES
LIQUID TRIM RELIEF VALVES
The liquid trim relief valve operates similarly to a conventional spring opposed pressure relief
valve and may be specified with a conventional design for general liquid applications or
specified with a bellows design for services where variable back pressure, high built-up back
pressure or corrosive fluids are present. The liquid trim design has a modified huddling
chamber design which provides smooth and stable valve operation on liquid service
applications and full lift and flow capacity at 10% overpressure.
3.6.1
Operating Characteristics
When system operating pressure reaches the specified set pressure, a small steady
stream of fluid begins to flow from the valve. Valve disk lift at this pressure is minimal.
As the system operating pressure increases into the range of 3% to 5% overpressure,
the valve then opens with a “pop” type lift towards the full lift position. At 10%
overpressure, the liquid trim valve is at full lift and rated capacity. As the operating
pressure begins to decay, flow through the valve decreases until the valve reseats with
a clean positive closing action at about 20% of rated flow.
The basic reason of the valve’s stability is that the huddling chamber is designed to
amplify hydraulic forces and that the blowdown is longer than for a conventional valve in
vapor service (nominally 10% instead of 7%) and longer than for a conventional valve in
liquid service, which has effectively zero blowdown. The presence of a blowdown gives
the valve an operating range rather than just an operating point. This is particularly
important in liquid systems which have very low capacitance as compared to vapor
systems. (Refer to Figure 3.2 B and C)
3.6.2
Applications
Liquid trim relief valves are required when a relief case exists where the relief fluid
remains a liquid throughout the valve and will not flash or vaporize across the valve
nozzle. If significant flashing or vaporizing is expected to occur, safety relief valves
should usually be used. If only small amounts of vapor (less than 10 volume percent)
are expected at the nozzle throat, liquid trim should usually be used as the fluid may
exhibit only slight compressible fluid characteristics.
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Provided that their operating characteristics are acceptable for a particular application,
liquid trim relief valves may be used for services which have relief cases where the relief
stream is expected to be all liquid, but which also have relief cases in which the relief
stream is all vapor or vapor/liquid. However, the potential effects of extended blowdown
for the vapor or vapor/liquid relief cases must be considered. Final selection should not
be made prior to consulting with the manufacturer.
3.6.3
Design Considerations
Design considerations for liquid trim relief valves are similar to those for conventional
spring opposed safety relief valves and balanced bellows safety relief valves. The
extended blowdown of a liquid trim relief valve may not be suitable for a combined
service application where the operating pressure exceeds the liquid trim relief valve’s
blowdown pressure. Also, some manufacturers have not certified their valves for vapor
service.
3.7
SPECIAL FEATURES
Pressure relief valves are available with a number of features, some of which may be
mandated by ASME code. Common examples are described below.
o
Lifting Levers
A lifting lever is a device which allows a pressure relief valve to be tested for proper
operation where corrosion or deposits could prevent normal operation of the valve or to
remove particles trapped under the seat as the valve closes. Code requires that lifting
levers be capable of opening the valve when the inlet pressure is 75% or more of set
pressure. ASME Section I requires lifting levers for all safety valves. ASME Section
VIII requires a lifting lever on air, steam, and hot water services above 140 °F (60 °C).
Lifting levers are rarely specified unless required by code.
o
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The purpose of the gag is to hold a pressure relief valve closed while equipment is
being subjected to pressure that is greater than the set pressure of the valve, as is the
case during a hydrostatic test. In refinery services, these are seldom specified.
o
O-ring (Soft) Seats
O-ring seats are used for difficult services (such as hydrogen, corrosives, nozzle icing,
etc.) or to provide positive closure at service pressures closer to the set pressure than is
possible with metal-to metal seats. Caution is required when specifying soft seats to
insure compatibility with the process fluid and temperature requirements.
o
Bolted Cap
A bolted cap is required whenever a pressure relief valve is provided with a packed
lever assembly. In other services a threaded cap is provided.
o
Spindle Extension
Spindle extensions are available for some models of spring opposed relief valves to
allow mounting of hydraulic lifting test devices or position transmitters. These are
seldom specified.
o
Weather Shield
Weather shields are available for Section I valves or open yoke Section VIII valves to
protect exposed springs in outdoor service.
o
Air Operated Lifting Device
This type of power actuator provides remote operation of the lifting lever. Usually these
consist of a pneumatic power piston which is mechanically connected to a lifting lever
assembly. A solenoid valve routes air to the piston when the valve is to be remotely
opened. These are occasionally used for Section I valves in the electric utility industry,
but never used in refinery applications. ASME Section I does contain a description of
when these types of devices are allowed.
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o
RELIEVING DEVICES
Open Yokes (Optional for Section VIII Only)
Some section VIII safety relief valves are available in open yoke construction instead of
the normal closed bonnet construction. This option is generally only used in steam
applications or high temperature applications when it is desirable to keep the spring as
cool as possible and the venting which will occur around the valve spindle is acceptable.
Yukong’s practice is not to use an open yoke construction unless the manufacturer
requires it due to service conditions. Generally all Section VIII steam valves are
specified with the standard closed bonnet.
o
Steam Trim
Some manufacturers offer a special steam trim for safety relief valves in Section VIII
steam service. This option usually consists of special thermally stabilized disk designs,
and special disk holder designs which allow for extra seating force on the disk. Not all
manufacturers offer such a trim, generally because they believe their standard designs
do not require special steam features. All steam services should have the suitability of
the valve trim reviewed with the manufacturer before purchase.
3.8
PILOT OPERATED PRESSURE RELIEF VALVES
Pilot operated pressure relief valves are considerably different than spring opposed pressure
relief valves. Pilot operated pressure relief valves do not have a large spring which opposes
the force of the inlet pressure. Instead a pilot, which, in simple terms, is a small spring
opposed pressure relief valve, routes process fluid to a dome above the valve disk. The
process pressure in the valve dome provides the force to hold the valve closed. When set
pressure is exceeded, the pilot vents the dome fluid and allows the process pressure under the
disk to force the valve to open.
This section describes how pilot operated pressure relief valves are constructed and operate.
There are additional descriptions of pilot operated pressure relief valves available with which
the engineer should be familiar. A general description appears in API RP 520, Part I.
Manufacturers’ catalogs provide detailed descriptions of how their various valves operate.
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3.8.1
RELIEVING DEVICES
Operating Characteristics
There are two basic types of pilot operated pressure relief valves, a piston type and a
diaphragm type.
3.8.1.1 Piston Type
The piston type pilot operated pressure relief valve, shown in Figure 3.4, has a
free-floating piston which acts as disk holder. Under normal operation, the
system operating pressure acts on the main valve seat at the bottom of the
piston and, by means of the pilot supply line (integral or remote from valve
body), process pressure is also applied to the pilot valve seat and the top of
the piston. There is an elastomer seal which prevents leakage from the dome
to the valve outlet.
The top of the piston is larger than the bottom seat area when the main valve
is closed, so there is a net downward force holding the piston on the nozzle.
Under static conditions the net downward seating force increases as the
pressure increases and approaches the set point. This is the opposite of the
conventional spring opposed pressure relief valve which reduces the net force
on the seat as set point is approached, which can result in simmering.
When the set pressure is reached, the pilot opens and partially or totally
(depending on pilot type) depressures the dome, thus reducing the force on
the top of the piston to the point where the upward force on the main valve
seat can overcome the downward loading. This causes a lifting of the piston
and the resulting flow through the main valve.
When the predetermined system blowdown pressure is reached, the pilot valve
closes, the full system pressure is diverted to the dome, and the piston moves
downward to close the main valve.
The standard pressure sensing point typically is in the main valve inlet nozzle
and assures that the pilot senses the pressure at the valve inlet whether the
main valve is closed or open and flowing. Integral sensing lines have total
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pressure sensors which measure stagnation pressure rather than static
pressure. However, the pilot sensing line may be located elsewhere if
necessary to provide sensing at another point in the system or to overcome
the effects of high inlet losses.
3.8.1.2
Diaphragm Type
The diaphragm type pilot operated pressure relief valve, shown in Figure 3.5 is
similar to a piston type except the piston is replaced by a flexible diaphragm
and disk. Usually these valves are used in low pressure, high capacity
applications. The diaphragm provides the seating force function of the piston.
The disk, which normally closes the main valve inlet, is attached to the flexible
diaphragm.
The external pilot valve serves the function of sensing process pressure,
venting the top of the diaphragm at set pressure, and reloading the diaphragm
once the operating pressure returns to normal. Similar to the piston type, the
seating force increases proportionally with operating pressure due to the
differential exposed area of the diaphragm.
3.8.1.3
Pilot Operating Description
A pilot operated pressure relief valve consists of two principal parts, a main
valve and a pilot valve. The pilot effectively is a small spring opposed
pressure relief valve. At set pressure, the pilot valve opens, and vents the
main valve dome area, reducing the pressure on top of the piston or
diaphragm, which allows the main valve to open. Likewise, when operating
pressure is reduced, the pilot valve closes at a predetermined blowdown, once
again loading the top of the piston or diaphragm and closing the main valve.
Pilot operated pressure relief valves operate with a precise action when set
pressure is reached and their operation is usually not affected by variations in
back pressure. However, some pilots may not be of a balanced design and
require venting to atmosphere.
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3.8.1.4
Pop and Modulating Action Pilots
Manufacturers typically offer pilots with either modulating or pop action. Pop
action pilot valves are characterized by a fully open main valve whenever the
sensed pressure exceeds set pressure and a fully closed main valve with a 5%
to 7% blowdown. In this sense, their operation is similar to a spring opposed
pressure relief valve. However, unlike spring opposed valves, pilot blowdown
can be easily set and tested. Pop action pilots can operate at higher maximum
system pressures than modulating valves, with an upper limit of approximately
5690 psig (392 barg, 400 kg/cm2 G).
Modulating type pilot operation is similar to a back pressure regulator. The
pilot opens the main valve only enough to relieve the overpressure, and
exhibits essentially 0% blowdown. This type of valve gives the maximum
conservation of process fluid and the least cost per relieving cycle. However,
the main advantages of the modulating pilot are enhanced stability and an
insensitivity to oversizing. Some types of modulating pilots also will operate
with stability at up to 30% inlet losses.
Generally, the trend is towards the use of modulating pilots due to their
enhanced performance characteristics. However, the engineer specifying the
pilot type must use caution in services such as auto refrigeration (i.e. LPG) or
services where the piston may become clogged at the seat during relief of
small loads. In these cases, pop action pilots should be used. Modulating
pilots also are not generally available for high set pressures [1000 psig (69
2
barg, 70 kg/cm G)].
3.8.1.5
Flowing and Non-flowing Pilots
Pilots are available in either flowing or non-flowing configurations. Non-flowing
pilots can often be used in the majority of applications and are preferred over
flowing pilots.
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Non-flowing pilots, with proper filtering, can be used in wet and/or dirty
services. Most non-flowing pilots are not balanced designs and need to be
vented to atmosphere.
3.8.1.6
Restricted Lift
Many pilot operated pressure relief valves are not produced with discrete
nozzle sizes as spring opposed valves are. Many times, a single nozzle may
be used for several different effective orifice sizes. The effective orifice size
required is obtained by installing lift restrictors in the valve dome. The lift
restrictor prevents the piston from reaching full lift and makes the curtain area
between the nozzle and disk the controlling flow area. Also, many pilot
operated valves are of a “semi-nozzle” design. This type of nozzle usually has
a lower coefficient of discharge than do the full nozzles found in spring
opposed valves.
3.8.2
Applications
Pilot-operated pressure relief affords the following advantages:
o
A pilot-operated valve is capable of operation at close to the set point and
remains closed without simmer until the inlet pressure reaches the set pressure.
It is possible to take advantage of this by reducing the normal 10% margin
between operating and set pressures, thus reducing vessel wall thickness
requirements.
o
Once the set pressure is reached, the valve opens fully and remains open, so
long as the set point is exceeded. As a result, it is not subject to chattering at
low discharge rates.
o
By locating the pilot valve pressure tap directly on the vessel being protected
(upstream of any inlet piping restrictions) a pilot-operated valve is less subject to
the chattering which is normally associated with high pressure drop inlet piping.
The capacity of the valve may have to be compensated for reduced inlet
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pressure. Remote sensing line sizing and routing should be reviewed with the
manufacturer.
o
When the pilot exhausts to the atmosphere, a pilot-operated pressure relief
valve is fully balanced like the balanced bellows valve. Therefore, its opening
pressure is unaffected by back pressure, and high built-up back pressure does
not result in chattering.
o
Pilot-operated valves may be satisfactorily used in vapor or liquid services up to
a maximum back pressure (superimposed plus built-up) of 50% of set pressure,
provided that the back pressure is incorporated into the sizing calculation. At
higher back pressures, capacity becomes increasingly sensitive to small
changes in back pressure. Back pressure up to 75% of set pressure may be
used, provided that this disadvantage is recognized.
o
With a simple test connection, both the pilot pop and reseat pressures can be
checked while the valve is in service.
o
A pilot-operated valve is sufficiently positive in action to be used as a
depressuring device. By using a hand valve, a control valve or a solenoid valve
to exhaust the piston chamber, one can open the pilot-operated valve and close
it at pressures below its set point from any remote location, without affecting its
operation as a pressure relief valve.
o
Modulating pilot-operated pressure relief valves exhibit very little blowdown.
This is an advantage for main gas pipeline and pressure storage applications,
where the narrow range of pressure cycling minimizes product losses resulting
from a release.
o
For applications involving unusually high superimposed back pressure, a pilotoperated valve may be the only possible balanced valve that is commercially
available, because of the mechanical limitations which apply to bellows.
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3.8.3
RELIEVING DEVICES
Design Considerations
When they were first introduced, the flexibility of pilot operated pressure relief valves
made them very attractive. However, especially when the standard pilot was a flowing
type, there were a lot of misapplications of the valves. Pilots became plugged with
particulates from the relief stream or in some cases, froze up due to pressure drop
across the pilot. Pilot operated valves in anything except extremely clean services got a
bad reputation.
Newer non-flowing pilot designs are not nearly as susceptible to problems, and they
have even greater functional flexibility, which makes them even more attractive as
compared to spring opposed pressure relief valves. However, they do not yet have the
versatility of application that spring opposed valves have, thus they cannot be
considered to be a universal replacement. The limitations of pilot operated valves still
must be recognized when they are being considered. Some limiting factors are as
follows:
o
Many pilot operated pressure relief valves are suitable for liquid services, but a
critical review with the manufacturer should be made before a pilot operated
valve is applied. The low capacitance of liquid systems makes valve stability a
concern. Pressure spikes or transients may cause unwarranted main valve
opening, sometimes at inlet pressures below the set pressure. Snubbers may
have to be installed in the pilot sensing line.
o
Pilot operated pressure relief valves have often had problems because their
installation and maintenance needs have not been well understood. Adequate
installation detail drawings and installation inspections are necessary for these
valves. Technicians need to be trained in pilot assembly and maintenance.
Most pilot operated pressure relief valves also require routine replacement of the
elastomers and plastics used in the pilot and main valve.
3.8.4
Special Features
o
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All pilots have small internal filters in their sensing ports. However dirty services
with catalyst, sand, rust, scale, or foreign particulates in the process require an
external pilot filter to maintain cleanliness within the pilot. The filter should be
equipped with a bleeder valve to purge the filtered material while the pressure
relief valve is in service.
o
Test Ports
The field test option consists of a field test valve and shuttle type check valve in
the pilot body. The field test option along with a pressure source may be used to
set the pilot set pressure in place on the main valve without requiring a system
over pressure and without interfering with the automatic pressure relief function
of the valve. This option can also function as a manual blowdown valve,
operating the main valve while the process fluid is below set pressure. The
design of field test systems should be reviewed with the manufacturer to verify
that the basic relief functions of the valve are not bypassed during testing.
o
Back Flow Preventers
In some pilot operated pressure relief valve designs, subjecting the discharge
side of the valve to a pressure that is higher than the inlet side can cause the
valve piston to lift even if the inlet pressure is less than set pressure. An
example in which this might happen is a pilot operated pressure relief valve in a
service with a low operating pressure, connected to a closed system which could
reach pressures higher than the valve’s normal inlet pressure. The same thing
could happen if a pilot operated relief valve is connected to a closed header
system, but the vessel it is mounted on is out of service.
To prevent unwanted opening of pilot operated pressure relief valves,
manufacturers offer optional backflow preventers. These devices are installed in
the vent line between the pilot and the valve outlet and maintain enough force on
the piston to keep it closed at all times that the inlet pressure is less than set
pressure.
o
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Snubbers
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In liquid services, hydraulic transients may cause unneeded opening of the main
valve. To prevent this, manufacturers offer hydraulic snubbers for installation in
the pilot inlet lines.
3.9
RUPTURE DISKS
3.9.1
Operating Characteristics
A rupture disk is a thin diaphragm, usually made of metal, mounted between flanges,
designed to burst at a designated differential pressure. It serves as a deliberately weak
element to protect vessels and piping systems against excessive pressure. Unlike
pressure relief valves, rupture disks do not reclose.
There currently are six major types of rupture disks available. Each of these is
described below. However, the generally preferred disk types are the prescored tension
loaded or prescored reverse buckling designs because of their higher allowed operating
pressures, non-fragmentation and low maintenance characteristics.
o
Conventional Tension Loaded Disks
Conventional tension loaded rupture disks, shown in Figure 3.6, are solid metal
disks that are normally prebulged by the manufacturer. This type of disk is
installed with the dome facing away from the process system being protected.
Accordingly, the disk metal is under tension. As system pressure approaches
the disk burst pressure, the dome stretches and then bursts due to excessive
tension loading.
Depending on disk size, burst pressure and temperature, a conventional disk
might be very thin. In such cases, it must be provided with support rings that
prevent disk crimping in the holder assembly. Vacuum supports may also be
required if the system pressure can be lower than the pressure on the
downstream side of the disk. In some cases, the support ring can act as the
vacuum support. To assure the use of the correct support rings and vacuum
supports, these are permanently attached to the disk.
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A conventional rupture disk will provide a good service life if the maximum
operating pressure in the system is at or below 70% of the disk burst pressure.
These disks can be used in gas service, in liquid-filled system, or in partially
liquid-filled system. These disks are the least expensive, and have the most
application restrictions, of the six types.
o
Prescored Tension Loaded Disks
Prescored tension loaded rupture disks, shown in Figure 3.7, are prebulged
metal disks which have been scored by the manufacturer to weaken them. They
avoid key drawbacks of a conventional tension loaded rupture disk;
fragmentation and the need for vacuum supports. The prescored pattern
ensures a predictable, nonfragmenting rupture. The disk can be fabricated so
that vacuum supports are not required. Typical maximum operating pressures
are up to 85% of burst pressure.
o
Composite Disks
Composite rupture disks, shown in Figure 3.8, are similar to conventional
tension-loaded rupture disks in their method of rupture, but they consist of
multiple, prebulged sections. Typically, a composite disk has a slotted metal
top section that defines the burst pressure by the size and location of its slots
and perforations. Under this pressure section is a plastic or metal membrane
that seals and isolates the pressure section from the process fluid. As with the
conventional disk, composite disks may require support rings and a vacuum
support.
Because of the composite construction assembly, composite disks allow the use
of corrosion resistant materials in lower-pressure services and smaller sizes than
is possible with a solid metal disk. In addition, a composite disk is less sensitive
than a conventional disk to operating pressure. Typical maximum operating
pressures for these disks are 80% of burst pressure.
o
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Reverse Buckling Disks with Knives
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Reverse buckling rupture disks function in a completely different way than a
conventional disk. These prebulged metal disks are installed with the dome
facing into the protected system, so the system pressure puts the dome into
compression. In contrast to the situation with a conventional tension-loaded
disk, the reverse-buckling dome does not stretch as the pressure approaches
the burst pressure. Instead a dent appears, which increases in size and causes
the disk to snap in the direction of lower pressure. In the knife blade design,
shown in Figure 3.9, the knife blades on the downstream side of the assembly
penetrate the disk as it snaps back, cutting it into three or more petals without
fragmentation.
Reverse-buckling disks can be made thick enough that no vacuum support is
needed, and can be used with operating pressure as high as 90% of burst
pressure. The knife blade design, while effective in ensuring that the burst is
non-fragmenting, introduces other significant concerns. Since the blades are
vital to the proper operation of the disk, any damage to them will adversely affect
the performance of the disk. Knife blades can corrode and get dull with time.
Dull blades may not properly cut the disk, resulting in elevated burst pressures
or reduced capacity.
Liquid filled or partially liquid-filled systems cause special problems for a
reverse-buckling disk with knives, and this arrangement should not be used on
those systems. Since liquids are essentially incompressible, reversal of the disk
occurs without significant snap action. The disk then acts like a damaged
reverse-buckling disk where the knife blades may fail to open the disk enough to
provide relief.
Reverse-buckling rupture disks with knife blades come in a wide range of sizes,
materials, and burst pressures and temperatures. These disks are also
available with coatings and linings to improve performance in corrosive services.
However, because of the requirement for the knife blades to always be sharp,
these types of disks should not be specified unless another type of disk cannot
meet the service requirements.
o
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Prescored Reverse Buckling Disks
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Prescored reverse buckling disks, shown in Figure 3.10, are prebulged solid
metal disks that have been weakened by the manufacturer along specified score
lines. They offer most of the benefits of knife blade reverse-buckling designs,
such as the lack of need for vacuum supports and permissible operating
pressure of up to 90% of burst pressure and avoid the potential problems that
knife blade designs have. Another advantage of these types of disks is that they
are often readily available with a zero manufacturing range.
The mechanism by which these disks snap back from their original position is
the same as for the knife blade design. But once snap-back has taken place,
the weakened score lines cannot withstand the combination of snapping force
and system pressure, so the disk ruptures along the score pattern. The score
pattern eliminates the possibility of fragmentation.
Because the snap-back of the disk is critical to proper functioning, these disks
should only be used in compressible services. As described for reverse-buckling
knife blade designs, they are subject to gradual reversal, which can result in the
disk failing to burst at its stamped pressure.
Prescored reverse-buckling rupture disks are available in the most common
sizes and materials for a limited number of burst pressures and temperatures.
Their operating characteristics and more inherent reliability make them the most
desirable for refinery applications.
o
Reverse Buckling Disk for Liquid Service
A reverse buckling disk design is also available for use in liquid and/or gas
service. This disk is designed to provide full opening in non-compressible liquid
service by utilizing a hinged design in combination with a circular score. The
hinge assures full opening along the score line, even if the reversal occurs
slowly. The hinge also retains the petal after reversal, thus preventing
fragmentation. This type of disk can be used with operating pressures up to
90% of burst pressure which provides an advantage over use of a conventional
disk in liquid service.
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Graphite Disks
In severely corrosive services, metal or composite rupture disks may not be
suitable. For these services, rupture disks constructed of graphite impregnated
with a binder material are available. A typical graphite disk is shown in Figure
3.11.
Graphite disks should only be used in services where the normal operating
pressure does not exceed 70% of the burst pressure. Graphite disks are
constructed of a brittle material and will fragment when they burst, so they are
not suitable for use under pressure relief valves. Graphite disks are primarily
used in the chemical industries where severely corrosive streams exist. There
are few such applications found in the refining industry.
3.9.2
Applications
Although well suited as primary pressure relief devices, rupture disks are generally not
used as primary devices in the refining industry, because rupture disks do not reclose.
A burst disk must be replaced before the process can be operated again. Rupture disks
are installed most frequently for the following reasons:
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To relieve an explosion pressure.
o
To avoid corrosion or plugging problems with a pressure relief valve.
o
To prevent fugitive emissions from a downstream pressure relief valve.
o
To handle high viscosity liquids and slurries.
o
To provide high volume capacity for an unusual condition in parallel with a tower
capacity relief valve sized for more typical upset conditions.
o
To provide rapid depressuring when required, in addition to preventing
overpressure.
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Care must be exercise when selecting and specifying rupture disk assemblies.
An analysis of the process conditions to confirm operating pressure, maximum
operating temperature, equipment design pressures, whether service is
incompressible or pulsating, and metallurgy requirements is crucial.
3.9.3
Design Considerations
The user must consider the following items when specifying a rupture disk:
o
The characteristics of rupture disks make specification of the burst pressure for
rupture disks more complicated than specifying a set pressure for pressure relief
valves.
o
The entire contents of the protected system are lost when the disk ruptures.
This necessitates a shutdown of the operation for replacement of the disk,
unless a block valve is installed upstream. In the latter case a CSO valve must
be used.
o
The actual bursting pressure may deviate by ± 5% of the designated bursting
pressure in the “as new” condition. (See Figure 3.17. In addition, a
manufacturer’s tolerance must be provided within which the disks will actually be
stamped). The narrower this allowance, the more expensive the disks will be.
The effect of fatigue in service may result in premature failure at lower
pressures. Therefore an added margin between operating and set pressure
equal to 20% must be provided. For this reason a rupture disk may require a
higher than normal equipment design pressure.
o
A rupture disk is sometimes installed upstream of a pressure relief valve, to
avoid leakage of high-cost materials, or to minimize corrosion or fouling of the
valve. Some means of detecting and relieving pressure buildup between the
disk and valve must be provided due to leakage through the disk as a result of
corrosion or some other cause.
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Minimum metal thickness makes fabrication of disks with very low burst pressure
impractical and unreliable. Generally the minimum available burst pressure is
around 20 psig (1.4 barg, 1.4 kg/cm2 G). For large disks (6-inch and larger)
made from aluminum, burst pressures as low as 15 psig (1.03 barg, 1.05 kg/cm2
G) may be available.
o
Rupture disks are differential devices. Any back pressure which exists will
directly affect the inlet pressure at which the disk bursts. If a back pressure
exists, it must be clearly specified and the specified burst pressure adjusted
accordingly. Consequently, stand alone rupture disks are not suitable for
variable back pressure situations unless the variation in back pressure is a very
low percentage of burst pressure.
o
The burst pressure of a rupture disk is a function of the temperature when it
bursts. The burst pressure of any rupture disk generally varies inversely with
respect to the coincident temperature. The sensitivity of the burst pressure to
temperature depends on the disk design and material. Of the common
materials, Inconel is the least sensitive to temperature variations, and aluminum
is generally the most sensitive. Due to their design, reverse-buckling disks are
only about half as sensitive to temperature as conventional disks.
o
In most instances the disk is installed at the end of a piping section where there
is normally no flow. Unless the system is well insulated, the normal disk
temperature can approach ambient temperature.
o
When used in combination with a safety relief valve, the capacity of that valve
must be derated by a factor of 0.9 or a value determined by testing the
combination. Inlet loss to the safety relief valve must still be evaluated including
the resistance of the rupture disks.
o
To ensure the safety of the system being protected, rupture disks must be failsafe devices. This has generally been accepted to mean that its design must
limit the burst pressure when the disk is damaged or improperly installed to less
than 1.5 times the system’s design pressure (MAWP). Under those
circumstances, any damaged or improperly installed rupture disk will, at worst,
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allow the system to be subjected to an overpressure equal to the hydrotest
pressure. Since this pressure has already been shown to be safe, the potential
for catastrophic failure of the system is minimized.
o
The engineer specifying the rupture disk should verify that the disk is
guaranteed to burst below 1.5 times stamped burst pressure when installed
upside down or damaged. Generally, the only types of disk that fall into this
category are the prescored tension loaded type and the prescored reverse
buckling type.
3.9.4
Rupture Disk Burst Pressure Example
Due to the importance of correctly specifying a rupture disk, a brief example of the
selection and specification process in provided:
As example of a specified burst pressure is a scenario where an ASME coded pressure
vessel has a MAWP of 150 psig (10.3 barg, 10.5 kg/cm2 G) and a maximum operating
pressure of 130 psig (8.9 barg, 9.1 kg/cm2 G). The operating pressure as a percentage
of MAWP is therefore 87%.
A prescored reverse buckling rupture disk is therefore selected which has a maximum
operating pressure factor of 90%. The least expensive standard manufacturing range
listed by the manufacturer is 10% for this type of disk. However, a burst pressure of
2
150 psig (10.3 barg, 10.5 kg/cm G) (full MAWP) and the coincident temperature and a
manufacturing range of +0% to -10% is provided. The stamped burst pressure, as
2
determined from burst test of the lot, may fall between 135 psig (9.1 barg, 9.5 kg/cm G)
2
and 150 psig (10.3 barg, 10.5 kg/cm G). Any disk received with a stamped burst
pressure less than 144 psig (9.9 barg, 10.1 kg/cm2 G), however, would not be
acceptable as the rupture disk maximum operating pressure factor of 90% would be
exceeded with a 130 psig (8.9 barg, 9.1 kg/cm2 G) system operating pressure. To
insure that the delivered rupture disk would be acceptable, a special manufacturing
range of +0% to -3% must be specified. All disks would then be stamped to burst
2
2
between 145 psig (10.03 barg, 10.2 kg/cm G) and 150 psig (10.3 barg, 10.5 kg/cm G).
3.9.5
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Holders
Rupture disk holders, also known as safety heads, come in a variety of types
and materials. Usually the minimum holder material should be stainless steel.
Typically a rupture disk holder is a two-piece unit consisting of a base flange
(inlet) and a holddown flange (outlet). The faces of these holders are machined
to grip a specific type of rupture disk and, together with the disk, form a leak
tight seal. Rupture disk holders come in bolted, insert, union and threaded
configurations. In most applications, insert type holders are used. The holders
are available with different features depending on the application. Knife blades
are installed for some reverse buckling type disks. Holddown flanges are
available with eyebolts, for attaching a hoist if desired, baffle plates for free
vents, and gauge taps for installing telltale indicators or drains.
A holder must be used with the model of rupture disk for which it was designed.
If the manufacturer of a disk is changed, the holder must also be replaced. If
the type of disk is changed, the manufacturer must be consulted to determine if
the holder must also be changed.
o
Vacuum Supports
Vacuum supports are typically required for conventional or composite tension
loaded disks. The main purpose of the vacuum support is to protect the rupture
disk from collapsing (failing in the reverse direction) during a vacuum or back
pressure cycle. Most vacuum supports are designed to withstand 15 psig (1.03
2
barg, 1.05 kg/cm G) back pressure and can be supplied for higher back
pressure on request. The vacuum support is designed for a specific disk and
should be permanently attached by the manufacturer. The vacuum support is
designed to open with the rupture disk to assure that there is no restriction to the
pressure relief area. The vacuum support can also serve as a protective ring
underneath the rupture disk to protect the seating area from dirt and debris and
to provide rigidity. For high vacuum applications, the assembly can be provided
with a soft ring seat which may be greased to provide additional tightness when
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high vacuum is encountered. The actual designs of vacuum supports vary with
the manufacturers.
o
Telltale Devices
A telltale devise is used when a rupture disk and a safety relief valve are used in
series. The ASME code states that the space between a rupture disk device
and a safety or safety relief valve shall be provided with a pressure gauge, a try
cock, free vent, or suitable telltale indicator. This arrangement permits detection
of disk rupture or leakage. Typically, the telltale indicator consists of a pipe
nipple, tee, an excess flow check valve, and a pressure gauge as shown in
Figure 3.12. The outlet side of the excess flow valve is threaded to permit a
discharge line to be installed if free venting is not desired and allows for routing
the product to a safe location. Under normal conditions, any leakage through
the disk is relieved through the excess flow valve. If the leakage is large, the
excess flow valve is forced closed and the pressure gauge will record a positive
pressure. If the rupture disk bursts because of overpressure, the excess flow
valve will close to prevent the release of material.
o
Burst Alarms
Another available method for telltale indication, is to use a pressure switch in
place of a pressure gauge. If pressure builds above the disk cavity, it trips an
alarm device at the site or at a remote location. Other remote alarm systems
are also available utilizing a normally open contact installed inside the burst
cavity which provides a switch closure upon rupture and also a system where a
wire is broken upon rupture to provide a normally closed contact arrangement.
Infrared photoelectric sensing is also being used to provide a positive response
to the bursting of a rupture disk. The use of electrical contact or photoelectric
devices are not recommended, however, due to the fact that they are not reliable
in detection of small pinhole leaks. Use of any device other than a telltale device
with a pressure gauge or pressure switch also does not satisfy ASME Code
requirements.
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3.10
RELIEVING DEVICES
OTHER TYPES OF PRESSURE RELIEF DEVICES
There are a number of pressure relief devices used in services which are outside the scope of
the ASME Code. Most of these devices are designed for special purpose applications for a
specific type of equipment. Some of the devices which may be found in refinery applications
are described below.
3.10.1 Surface Condenser Pressure Relief Valves
In larger steam turbines, surface condensers are used to provide a turbine exhaust
pressure of less than atmospheric pressure. If cooling water is lost to the surface
condenser, or if the turbine discharge is blocked-in, the turbine and surface condenser
could be over pressured. In these non-ASME services, a special design low pressure
valve is used. These types of valves may also be applicable to installations which use a
combination of condensers and eductors to establish the turbine exhaust vacuum.
These valves usually do not contain any springs or other adjustment mechanism and
often have a weighted plate which sits on a valve seat. When the pressure in the
condenser exceeds the set pressure of the valve, the plate lifts and relieves the stream.
Surface condenser pressure relief valves usually are of a straight through design (not
an angle design as most relief valves are) and can be specified either for horizontal or
vertical installation. This must be specified at the time of purchase, as the valve
designs are different for each orientation.
Since they operate at a vacuum, surface condenser pressure relief valves generally
have a water seal on the valve seat. Usually they are designed to have a continuous
flow of water to the seals, with the seal level being maintained by an overflow weir. An
optional water seal gauge glass is available.
These valves are not governed by any design standard such as API Standard 526. The
engineer should be careful in reviewing specifications for these valves as the standard
materials may not be acceptable. In many of these valves, the standard body material
is cast iron, which is usually not acceptable in refinery applications. Carbon steel bodies
should be specified. The valve manufacturers also make considerable use of bronze
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internals. In refinery services, the preferred trim is stainless steel. Both of these
materials are non-standard, and may require longer delivery cycles.
These valves are not sized by techniques normally associated with pressure relief
valves. Their sizing is governed by NEMA Standards Publication No. SM23 and are
based on the steam loads. However, the user must check the actual turbine to ensure
that the valve is designed for maximum throttle load. Once a total steam load is
determined, the required valve size is determined from a capacity table provided by the
manufacturer. Because of the low differential pressure across these valves, they tend
to be quite large. The minimum valve size recommended is 6 inches (152 mm), with
valves of up to 36 inches (915 mm) available. The valves are sized for an inlet pressure
of 10 psig (0.7 barg, 0.7 kg/cm2 G).
Another consideration in sizing these valves is whether or not continued operation of the
turbine (at reduced power) is desired. If continued operation is desired, larger valves
are required to pass the flow at lower pressures. The Heat Exchange Institute
recommends that if this is desired, a separate gate or butterfly valve be provided for
non-condensing operation. However, in refinery applications, operation in a noncondensing mode is usually not practical.
Set pressures for the valve should be determined by consulting with the turbine
manufacturer and the surface condenser manufacturer. Normally the turbine will
govern the set pressure, especially if non-condensing operation is required. Set
pressure is usually fixed by the weight of the valve disk and is not adjustable.
3.10.2 Sentinel Valves
Sentinel valves are installed on non-condensing steam turbine cases to warn that the
case is being unacceptably back pressured or that a failure in the turbine wheel or
buckets has occurred. Sentinel valves are generally not ASME certified even though
2
they may have set pressures much higher than 15 psig (1.03 barg, 1.05 kg/cm G).
This is acceptable since they are not installed on ASME equipment and are not true
pressure relief valves because they are not intended to provide overpressure protection.
Sentinel valves are very small, usually with ½-inch to 1-inch inlets, and are normally
installed directly on the turbine case. Threaded inlet connections are used. As stated
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above, the main function of a sentinel valve is to warn of a potentially dangerous
situation. They are sized to pass only a nominal flow, and are designed to emit a loud
whistle when they are relieving.
3.10.3 Pressure/Vacuum Breather Valves
3.10.3.1
Introduction
Pressure vacuum breather valves are typically installed on low pressure fixed
roof storage tanks. They may be of a weighted pallet type, spring loaded
pallet type or pilot operated. Normally these valves are supplied by the tank
manufacturer.
Tanks and vessels designed for pressures of 15 psig (1.03 barg, 1.05 kg/cm2
G) and less are usually equipped with pressure-vacuum relief valves to meet
normal and emergency venting requirements. The types of vent devices
include pressure relief valves, vacuum relief valves, combination pressurevacuum valves, open vents, and hinged manhole or gage hatch covers.
3.10.3.2 Non-Refrigerated Tanks
Normal venting from non-refrigerated, fixed roof tanks is accomplished by
pressure relief valves, pressure-vacuum valves, or with open vents with or
without a flame arresting device in accordance with the following requirements:
o
Pressure relief valves are applicable on tanks operating above
atmospheric pressure. In cases where a vacuum can be created within a
tank, vacuum protection may be required.
o
Pressure-vacuum valves are recommended for use on atmospheric
storage tanks in which oil with a flash point below 100 °F (38 °C) is
stored and for use on tanks containing oil that is heated above the flash
point of the oil. A flame arrestor is not considered necessary for use in
conjunction with this type of valve.
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Open vents with a flame-arresting device may be used in place of
pressure-vacuum valves on tanks in which oil with a flash point below
100 °F (38 °C) is stored and on tanks containing oil that is heated above
the flash point of the oil.
o
Open vents without a flame-arresting device may be used for tanks in
which oil with a flash point of 100 °F (38 °C) or above is stored.
When tanks are exposed to fire, the venting rate may exceed the normal
venting rate. In such cases, emergency venting requirements are met by one
of the following means:
o
Larger or additional pressure relief valves, pressure-vacuum valves or
open vents as limited by rules for normal venting devices.
o
A gage hatch that permits the cover to lift under abnormal internal
pressure.
o
A manhole cover that permits the cover to lift when exposed to abnormal
internal pressure.
o
A connection between the roof and the shell that is weaker than the
weakest vertical joint in the shell or shell-to-bottom connection. Check
with the vessels engineer to determine if this condition is met.
Set and relief pressures are generally established as follows:
o
Atmospheric pressure tanks (API Std 650 tanks):
Pressure relief devices for normal venting have a set pressure which is
50 percent of the design pressure (gauge). The devices are fully open at
25 percent over set pressure (gauge).
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Pressure relief devices for emergency venting are set to open at twothirds of the design pressure (gauge) and are fully open at 20 percent
over set pressure (gauge).
o
Low pressure tanks (API Std 620 tanks)
Pressure relief devices for normal venting are fully open at the design
pressure which is 10 percent over the set pressure (gauge)
Pressure relief devices for emergency venting are set to open at the
design pressure and are fully open at 20 percent over set pressure
(gauge).
Vacuum relief devices for both types of tanks are typically set to open at 50
percent of the vacuum for which the tank is designed. The devices are also
generally designed to be fully open at the design vacuum.
Floating roof and lifter roof tanks are equipped with open vents for pressure
and vacuum protection. The vents are normally furnished by the tank vendor.
WARNING: The maximum filling and emptying rates of tanks are not limited to
transfer rates from other pumps. Where two or more tanks have a common
line to any destination, it is sometimes possible to inadvertently have both
tanks open to one another during valve realignment operations. The
maximum filling and emptying rates may occur when the liquid within the tank
with the higher level transfers into the tank with the lower level. This rate is a
maximum when one tank is full and the other empty. Failure to consider this
case has led to collapse of the emptying tank where the vacuum relief was
insufficient and rupture of the filling tank where the pressure relief was
inadequate.
3.10.3.3
Refrigerated Tanks [up to 15 psig (1.03 barg, 1.05 kg/cm2 G)]
Normal and emergency venting is accomplished by means of pressure relief
valves and vacuum relief valves or pressure-vacuum valves. These tanks are
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designed for pressures of 15 psig (1.03 barg, 1.05 kg/cm2 G), or less and a
vacuum of 1.5 inches of water [0.054 psig (0.0037 barg, 0.0038 kg/cm2 G)].
When tanks are exposed to fire, the emergency venting rate may exceed the
normal venting rate. In such cases, additional capacity is required for
emergency venting.
In general, pressure relief devices for normal venting are fully open at the
design pressure which is 10 percent over the set pressure (gauge). Pressure
relief devices for emergency venting are set to open at the design pressure
and are fully open at 20 percent over the set pressure.
Vacuum relief devices are set to open at 50 percent of the vacuum for which
the tank is designed. They are fully open at the design vacuum.
3.10.4
Explosion Hatches
Most rupture disk manufacturers also offer large area, low pressure rupture devices,
typically referred to as explosion doors or explosion vents. These devices are typically
used in solids storage vessels such as grain elevators, dust collection systems and
some boilers and furnaces. They are also used in refinery applications.
For example, a vessel which operates at essentially atmospheric pressure and is
subject to internal explosion, such as an asphalt oxidizer, should be protected by an
explosion hatch. The hatch consists of a hinged metal cover fitted over an opening on
top of the vessel and sealed by its own weight. For vessels which normally operate at a
slight positive pressure, a tight seal is achieved by the use of hold-down brackets with
shear pins, rather than by increasing the weight of the hatch which would increase
inertia and prevent quick opening. One or more hatches may be provided for a single
vessel. If the need for such a device is identified, careful coordination with the
equipment manufacturer and the vent manufacturer is required.
3.10.5
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There are a number of pressure relief valves on the market that are not ASME
approved valves. These generally are found in the smaller sizes, but occasionally larger
ones are found. In refinery applications, there is little reason to purchase anything
except an ASME approved pressure relief valve, even in services where it is not
required.
3.10.6
Liquid Seals
In some cases, a hydraulic loop seal may be used for relieving overpressure on
equipment operating at pressures slightly above atmospheric. Examples are certain
naphtha fractionators with total condensation, where the seal would be installed on the
distillate drum vapor space, discharging to the atmosphere. The seal consists of a
simple U-tube containing a suitable liquid (normally water) with the seal depth and
diameter sized to pass the maximum relieving flow at the required design pressure.
The following design features should be incorporated:
o
Continuous water makeup and overflow on the seal loop, to ensure that the seal
is always made during normal operation, and reestablished after a blow.
o
Adequate winterizing, where necessary, to prevent freezing of the seal.
o
Safe disposal of the effluent seal water, considering possible contamination by
process fluids.
o
It must be acceptable to discharge the process fluid to atmosphere.
o
It must not be possible to discharge liquid hydrocarbons through the
atmospheric vent.
o
The vent line must be protected against flashback.
Although liquid seals are relatively simple, reliable, and inexpensive, they are of limited
application, because of the difficulty in meeting all of the criteria listed above. Also, they
may not be too practical where vacuum conditions are encountered.
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3.10.7
RELIEVING DEVICES
Vacuum Relief Valves
Normally, vessels and other equipment are designed to accommodate vacuum
conditions that may occur. Occasionally, vacuum relief devices are provided which
permit inflow of air, nitrogen or fuel gas when required. The risk of using air to relieve
vacuum should be considered prior to installing such systems. Information on the
selection, sizing and installation of vacuum relief valves may be obtained from
manufacturers’ literature.
3.11
PRESSURE RELIEF DEVICE SIZING
Pressure relief device sizing consists of computing the flow area required for the device to pass
the required relief load at relieving conditions. The methods and sizing equations which should
be used in determining pressure relief device sizes are described in this section. While the
methods described here will virtually always result in relief devices which are properly sized for
the specified relief conditions, final selection of relief devices should not be made until the
sizing has been reviewed with the manufacturer and checked against the manufacturer’s
published capacity data.
There are three basic ways to size pressure relief valves. One is to use the ASME rated data
for a specific device. The second is to use generic formulas such as those published by API,
and the third is to use manufacturer’s published equations, capacities and coefficients. Of the
three methods, use of the API equations is the most common, particularly for preliminary sizing
and specification and is discussed in this section.
Rupture disk sizing is dependent upon the type of application. Section 3.11.6 discusses sizing
of rupture disks used as primary relief devices and in combination with pressure relief valves.
3.11.1
API Sizing Equations
In order to use the ASME sizing equation, the engineer must know the actual nozzle
area and the rated coefficient of discharge for the pressure relief valve. These values
are different for every manufacturer and model of valve, so a specific pressure relief
valve must be selected in order to be sized using the ASME methods. This may be the
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case when final purchase specifications are being prepared, but during initial design an
engineer does not always know what manufacturer and model number of pressure relief
valve will be installed.
To address this problem, the API and the pressure relief valve manufacturers developed
a generic sizing equation which is suitable for most pressure relief valve sizing. Several
forms of this equation for use with vapors and gases, steam, and liquids are published
in API RP 520, Part I and are repeated in the following sections. The API equations are
based upon use of API “effective areas”, which allows a common basis for sizing and
selection of pressure relief valves without first having to know the actual nozzle area
and coefficient of discharge.
The API equations represent a consensus among the API members and the pressure
relief valve manufacturers to define generic equations that can be used to size pressure
relief valves. In building pressure relief valves to meet API Standard 526, the
manufacturer guarantees that a valve with a given API effective area will flow at least as
much as would be predicted by the API equation. The actual capacity will vary from
manufacturer to manufacturer, and may be slightly greater than the API flow rate.
Generally the actual rated capacity is in the range of 1% to 3% greater than would be
predicted from the API equations.
3.11.1.1
Vapor Equation - Critical Flow
This equation is the most used form of the API sizing equations and is similar
to the ASME equation, but differs in the values of the coefficients and
constants used. As with the ASME equation, critical flow across the valve
nozzle is assumed. The basic API vapor equation is:
A
A
A
=
=
[W/(CKdP1Kb)][(TZ)/MW]0.5
1.316[W/(CKdP1Kb)][(TZ)/MW]
3.1 (English)
0.5
3.1 (Metric)
0.5
=
1.342[W/(CKdP1Kb)][(TZ)/(MW)]
3.1 (Metric)
=
Required API effective nozzle area, in (cm )
Where:
A
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3.11.1.2
W
=
Relief flow, lb/hr (kg/hr)
C
=
Constant based upon heat capacity ratio of the gas
=
520{ k [2/(k+1)](k+1)/(k-1) } 0.5
k
=
heat capacity ratio of the vapor, Cp/Cv
Kd
=
API effective coefficient of discharge constant
value = 0.975
Kb
=
Back pressure correction for bellows valves (Figure 3.14)
P1
=
Inlet pressure, psia (bara, kg/cm2 A) (set pressure +
overpressure + atmospheric pressure)
T
=
Inlet temperature, °R (°K)
MW =
Gas molecular weight
Z
Compressibility factor at valve inlet
=
3.2
Effective Areas and Coefficient of Discharge
The API equation represents a “generic” pressure relief valve which has a
certain capacity for a given “effective” nozzle area. The effective areas are
defined by the API letter designations found in API/ANSI Standard 526, 4th
Edition, 1993, Tables 2 through 15. Table 3.1 at the end of this chapter lists
the API nozzle sizes and effective areas.
When the manufacturer constructs a valve to API Standard 526, the actual
orifice area is designed so it will flow at least the rate predicted by the API
equation for the effective area for which the valve will be designated. In
reality, the actual nozzle areas are always larger than the effective area and
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the actual coefficient of discharges are always lower than the 0.975 used in
the API equation. By ASME Code, the maximum rated coefficient used can
only be 0.90. Typical rated coefficients fall in the range of 0.85 to 0.87. Table
3.2 presents some typical values of effective and actual areas and coefficients
of discharge.
3.11.1.3
Back Pressure Effects
The ASME equations are intended to be used for pressure relief valve capacity
rating where no back pressure is imposed on the valve, and therefore do not
provide for any means of accounting for the effect of back pressure. In order
to size a pressure relief valve for a real application, some means of accounting
for the effects of back pressure need to be made. Back pressure has two
effects on a pressure relief valve:
o
It may reduce valve lift and capacity. This can affect performance to the
point where the pressure relief valve is unstable.
o
It may cause the valve nozzle to go sub-sonic at the throat.
Conventional safety relief valve capacity is reduced under sub-sonic
conditions, as shown in Figure 3.13, when the total constant backpressure
exceeds about 50% of the set pressure (absolute). However, valve instability
can occur, before capacity reduction in conventional valves, when the built-up
backpressure reaches about 10% of the spring (differential) set pressure. If
the total backpressure with a conventional safety relief valve is high enough to
cause the nozzle to go sub-sonic, but the built-up backpressure is not so large
that instability could occur, the methods described in Section 3.11.1.4 should
be used.
Testing of bellows pressure relief valves have shown that their capacities are
affected when back pressure reaches about 30% of set pressure. Therefore
the Kb coefficient was introduced to account for this effect. Figure 3.14 shows
how the Kb for bellows pressure relief valves varies with total back pressure.
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The back pressure used for sizing should be the sum of the constant back
pressure plus any built-up back pressure.
Figure 3.14 shows values of Kb for up to 50% back pressure. If a specific
application of a bellows pressure relief valve requires greater than 50% back
pressure, the valve manufacturer should be consulted to determine the effect
on capacity. Also, the values of Kb contained in Figure 3.14 are typical values
which do not apply to any specific manufacturer. The curves published by
manufacturers for their valves will generally present slightly larger values for
Kb.
3.11.1.4
Subcritical Vapor Flow
API RP 520, Part I, presents equations for sizing for sub-critical flow. This
equation is more complex than the critical flow equation, but is derived from
the same basic formula. Actually the critical flow equation is a simplified form
of the more general sub-critical equation.
A pressure relief valve should be sized based upon sub-critical flow whenever
the pressure ratio across the valve is less than the critical pressure ratio. This
critical pressure ratio is calculated by:
Rc = [(2)/(k+1)]
(k)/(k-1)
3.3
The sub-critical sizing equation is:
A = [W/(735 F2 Kd )][(Z T)/(MW P1 (P1-P2))]0.5
A = 1.316[W/(735 F2 Kd )][(Z T)/(MW P1 (P1-P2))]0.5
A = 1.342[W/(735 F2 Kd )][(Z T)/(MW P1 (P1-P2))]0.5
3.4 (English)
3.4 (Metric)
3.4 (Metric)
In which F2 is defined by:
F2 = {[(k)/(k-1)][r
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(k-1)/k
)/(1-r)]}
0.5
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Note: API RP 520 Part I, also provides graphs for determining F2 rather than
computing it analytically.
Where:
Rc
=
Critical pressure ratio
A
=
Required API effective nozzle area, in (cm )
W
=
Relief flow, lb/hr (kg/hr)
r
=
Ratio of outlet to inlet pressure = P2/P1
k
=
Heat Capacity ratio, Cp/Cv
Kd
=
API effective coefficient of discharge, constant
value = 0.975
P1
=
Inlet pressure, psia (bara, kg/cm2 A) (set pressure +
overpressure + atmospheric pressure)
P2
=
Outlet pressure, psia (bara, kg/cm A) (atmospheric pressure
+ back pressure)
T
=
Inlet temperature, °R (°K)
2
2
MW =
Gas molecular weight
Z
Compressibility factor at valve inlet
=
2
The above equations should only be used for conventional pressure relief valves
with sub-critical pressure drops. If bellows valves are used, the equation
presented in Section 3.11.1.1 should be used with the appropriate Kb factor.
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3.11.1.5
Steam Flow
API RP 520 publishes a sizing equation for steam which is similar to the ASME
equation. Critical flow is assumed by this equation. The equation is:
A = W s / (51.5 P1 Kd Kn Ksh )
A = 1.316 (W s) / (51.5 P1 Kd Kn Ksh )
3.6 (Metric)
A = 1.000 (Ws) / (51.5 P1 Kd Kn Ksh )
3.6 (Metric)
3.6 (English)
Where:
2
2
A
=
Required API effective nozzle area, in (cm )
Ws
=
Steam relief flow, lb/hr (kg/hr)
Kn
=
Napier correction factor. This value = 1.0 for steam
2
pressures at or below 1500 psig (103 barg, 105 kg/cm G).
See API RP 520 if above this pressure.
P1
=
Inlet pressure, psia (bara, kg/cm2) (set pressure +
overpressure + atmospheric pressure)
Ksh
=
Superheat steam correction factor. Obtain from Table 10
in API RP 520 Part I. For saturated steam, Ksh = 1.0
Kd
=
API effective coefficient of discharge, constant
value = 0.975
3.11.1.6
Liquid Flow - Liquid Trim Relief Valves
The API sizing formula for liquid trim relief valves flowing liquid is:
A = [Q/(38 Kd Kw Kv)][G/(P1 - P2)]
0.5
A = 7.456 [Q/(38 Kd Kw Kv)][G/(P1 - P2)]
0.5
A = 7.528 [Q/(38 Kd Kw Kv)][G/(P1 - P2 )]
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3.7 (English)
3.7 (Metric)
3.7 (Metric)
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Where:
3.11.1.7
A
=
Required API effective nozzle area, in2 (cm2)
Q
=
Liquid relief flow, gpm (m3/hr)
G
=
Specific gravity of liquid at flowing conditions
Kd
=
Effective coefficient of discharge obtained from the liquid
trim relief valve manufacturer. For preliminary sizing, a
value of 0.65 may be used.
Kw
=
Correction factor due to back pressure. This factor is
required only for bellows valves. A curve for Kw is
provided in Figure 3.15
Kv
=
Correction factor for liquid viscosity. A curve for Kv vs
Reynolds number is provided in API RP 520, Part I
(Figure 32). Kv is equal to 1.0 for Reynolds Numbers
above 40,000.
P1
=
Inlet pressure, psig (barg, kg/cm2 G) (set pressure +
overpressure)
P2
=
Outlet pressure, psig (barg, kg/cm G) (back pressure)
2
Liquid Flow - Conventional Pressure Relief Valves (Capacity certification not
required)
The API sizing formula for conventional pressure relief valves flowing liquid is:
0.5
3.8 (English)
A = [Q/(38 Kd Kw Kv Kp )][G/(1.25Ps - Pb )]
0.5
A = 7.456[Q/(38 Kd Kw Kv Kp )][G/(1.25Ps - Pb )]
3.8 (Metric)
A = 7.528[Q/(38 Kd Kw Kv Kp )][G/(1.25Ps - Pb )]
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Where:
A
=
Required API effective nozzle area, in2 (cm2)
Q
=
Liquid relief flow, gpm (m3/hr)
G
=
Specific gravity of liquid at flowing conditions
Kd
=
Effective coefficient of discharge obtained from the liquid
trim relief valve manufacturer. For preliminary sizing, a
value of 0.62 may be used.
Kw
=
Correction factor due to back pressure. This factor is
quired only for bellows valves. A curve for Kw is provided
Figure 3.15
Kv
=
Correction factor for liquid viscosity. A curve for Kv vs
Reynolds number is provided in API RP 520, Part I
(Figure 32). Kv is equal to 1.0 for Reynolds Numbers
above 40,000.
Kp
=
A correction factor for accumulation pressure of other than
25% above set pressures. A curve of Kp versus
accumulation is presented in Figure 3.16. For 10%
accumulation, Kp = 0.60, and for 21% accumulation,
Kp = 0.92.
3.11.2
Ps
=
Set pressure, psig (barg, kg/cm2 G)
Pb
=
Back pressure, psig (barg, kg/cm2 G)
Manufacturer’s Equations
Manufacturers publish equations which may be used to size their pressure relief valves.
Along with the equations, tables of nozzle areas, coefficients of discharge and other
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constants and correction factors may also be published. All of these equations are
based upon either the API or ASME equations, but seldom do the values inserted into
the equations directly correspond to ASME or API values. Each manufacturer seems to
have different methods of expressing nozzle areas and discharge coefficients, and
many times derating factors or differences from API effective areas and coefficients of
discharges are buried in constants or table factors.
Some of the ways that the manufacturers manipulate the basic sizing equations to
incorporate their coefficients include:
o
Combining C and Kb into one coefficient. Often this combined factor is
confusingly represented as C or K in a manufacturer’s formula.
o
Providing constants in their formulas which correct non-derated coefficients of
discharge or non-API areas in sometimes obscure ways. One manufacturer
uses an “effective” discharge coefficient and then uses a numerical constant to
correct the effective coefficient to the ASME rated coefficient.
o
Presenting two sets of equations, one for use with API effective areas and
discharge coefficients and a second for use with ASME coefficients.
This results in what can be a very confusing situation, especially when trying to compare
pressure relief valves of different manufacturers. Used correctly, all of the
manufacturer’s equations will properly size their own pressure relief valves. However,
the engineer must be careful to use the coefficients and constants defined by the
manufacturer and not mix them with API or ASME values. This can be a real challenge
when pseudo-API and pseudo-ASME coefficients are presented in the same catalog.
Fortunately, most manufacturers now have PC based computer programs available,
usually at no charge to customers, which will perform all sizing calculations and actually
recommend a pressure relief valve (from that manufacturer’s catalog, of course). A list
of programs available is contained in Appendix C. A final check of sizing using the
manufacturer’s programs or equations is recommended before a purchase specification
for a pressure relief valve is issued. When a calculated required area is just slightly
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larger than a standard orifice size, a check using the manufacturer’s equations or ASME
rated capacities may result in a smaller pressure relief valve size being required.
3.11.3
Pilot Operated Pressure Relief Valves
Pilot operated pressure relief valves may be sized in the same manner that spring
opposed pressure relief valves are sized. There are a few differences that should be
kept in mind.
o
API sizing equations may not be applicable to some pilot operated pressure
relief valve designs because of significantly lower coefficients of discharge in
those valves with semi-nozzle designs. Pilot operated pressure relief valves
should be sized using manufacturer’s equations or ASME equations with actual
coefficients.
o
Pilot operated pressure valves are not sensitive to back pressure, so values of
Kb for critical flow should always be taken to be 1.0. If the back pressure is so
high that the critical pressure ratio is not reached, a subsonic flow formula must
be applied.
o
Prior to final purchase, all pilot operated pressure relief valve sizing should be
checked using the manufacturer’s equations.
Even though API Standard 526 is not applicable to pilot operated pressure relief valves,
most pilot operated pressure relief valve manufacturers design their valves to be
consistent with API effective nozzle areas and API equations.
Many pilot operated pressure relief valves use a “semi-nozzle” design, which usually
has an actual coefficient of discharge which is lower than the full nozzle found in spring
opposed pressure relief valves. Some designs may also use oversized nozzles and
restricted lifts. This means that the actual area of the nozzle may be larger than the API
effective area by a greater margin than for a spring opposed valve. As long as
consistent coefficients and areas are used, this doesn’t present a problem. However,
the error caused by mixing of effective and actual values can be larger for pilot operated
valves than for spring opposed valves.
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3.11.4
RELIEVING DEVICES
Safety Valves
The API sizing equations are not intended for application to safety valves.
Manufacturers usually do not rate safety valves in terms of API effective areas (some
appear to do so, but close examination shows that this is only an approximation).
ASME Section I is quite specific when it comes to determining the steam rate that a
safety valve has to pass, and does not provide a safety valve sizing equation, such as
supplied by Section VIII in Appendix 11. The procedure for sizing a safety valve is to
determine the required steam relief capacity as required by Section I and select the
appropriate valve from a capacity table. Manufacturers do provide equations which can
be used to compute safety valve capacities at over pressures of other than 3%.
3.11.5
Sizing Procedures
Described below are general recommended procedures for sizing of pressure relief
valves for various applications. Special considerations for each type of pressure relief
valve commonly encountered are described in the separate sections below.
The basic steps of pressure relief valve sizing are:
o
Perform preliminary sizing using the API sizing equation applicable to the
situation. This will result in the correct valve size in almost all applications. If
there is an error in valve size, the API equations will usually require one nozzle
size larger than actually needed. Manufacturer’s equations should be used for
liquid trim or pilot operated pressure relief valves.
o
After selection of a manufacturer and model number, back-check the valve
sizing using the manufacturer’s sizing equations.
o
If the required area is very slightly above a standard orifice size (within less than
5% of standard area), check the capacity against the ASME rated capacities.
Most pressure relief valves have slightly more capacity than would be indicated
by the API sizing equations. This means that the equations will always give an
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area slightly larger than the area that is really needed to pass a given flow rate.
This is important to recognize when situations arise where a calculated area is
just large enough to require the next larger nozzle size. Generally, this occurs
when existing installations are being evaluated to identify the effects of changes
in an existing system. If this is encountered, the actual valve capacity should be
checked against the required capacity. Occasionally a larger orifice may not be
required.
3.11.5.1 Spring Opposed Safety Relief Valves - Vapor Service
Vapor or gas sizing for safety relief valves with conventional or bellows trim
should be calculated using the API equations shown in Section 3.11.1. Final
valve sizing must be verified against the manufacturer’s rated capacities prior
to purchase.
3.11.5.2
Spring Opposed Safety Relief Valves - Liquid Service/Conventional Trim
Conventional trim safety relief valves may not be used in new liquid services
unless the services are not covered by the ASME Code. A liquid trim pressure
relief valve should always be considered for use over a standard safety relief
valve for liquid service.
Where safety relief valves have been approved for use in liquid service, the
required orifice area may be calculated using the API equation in Section
3.11.1.7. Safety relief valves in non-ASME services should be sized with 10%
accumulation when used to protect pumps or vessels, and 20% accumulation
when used only to protect piping.
3.11.5.3
Spring Opposed Pressure Relief Valves - Liquid Service/Liquid Trim
The required orifice size for liquid trim relief valves may be determined using
manufacturer’s equations and coefficients or the equation shown in Section
3.11.1.6. It should be noted however, that the API equation may not be
applicable to all manufacturers. The engineer is once again cautioned that
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some manufacturers present equations and coefficients in terms of actual
orifice area rather than API equivalent orifice area.
Liquid trim relief valves are required for new ASME services where relief valve
sizing is governed by a case in which the relief fluid remains a liquid
throughout the valve and will not significantly flash or vaporize across the valve
nozzle. If the flashing or vaporization is expected to exceed 10 volume
percent, a safety relief valve should be used and sized for vapor/liquid service.
For combination services, the manufacturer should be consulted to determine
the proper selection of valve type.
3.11.5.4
Two Phase or Flashing Flow Service
Based on the API RP 520, Part I, Seventh Edition, new pressure relief valves
in services with two phase or flashing flow should be sized based on the HEM
(Homogeneous Equilibrium Model) by DIERS.
For existing installation, pressure relief valves in services with two phase or
flashing flow were sized based upon the method outlined in API RP 520, Part I,
earlier editions. This method (slightly modified) is as follows:
o
If the flow is flashing or if additional vaporization occurs across the
pressure relief valve nozzle, compute the vapor component of the stream
based upon an adiabatic, isentropic flash from inlet conditions to the
critical flow pressure ratio.
o
Compute the required area to pass the vapor load computed using either
API or manufacturer’s equations as described in Sections 3.11.1 and
3.11.2.
o
Compute the required area to pass the liquid load using either API or
manufacturer’s equations as described in Sections 3.11.1 and 3.11.2. In
applying these equations, if there is more than 10 volume percent vapor,
do not apply the derating factor of 0.6 commonly applied to safety relief
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valves in liquid service. Pressure drop for the liquid component of the
stream should be limited to the critical pressure drop of the vapor.
o
Select a safety relief valve which has an orifice area greater than the
sum of the areas required to pass the vapor component and the liquid
component of the two phase stream.
Some refiners have adopted a new method of calculating relief valve sizes for
two phase flow which is a version of the HEM (Homogeneous Equilibrium
Model) procedure proposed by DIERS. In the absence of a refiner’s method,
the HEM method as described in API RP 520, Section I, Seventh Edition,
should be used. This method results in larger calculated required areas for
relief valves.
3.11.5.5
Combination Services
Combined services are defined as services which have one or more relief
cases which are all liquid, but which also have other relief cases in which the
relief stream is all vapor or vapor/liquid. In these cases, the selection of a
pressure relief valve and orifice size is dependent upon relative area required
by each case. If the required area for the liquid relief case (based on liquid
trim equations) is 40% or greater than the required area for a vapor or
vapor/liquid (based on vapor formulas), liquid trim is normally selected.
However, the potential effects of extended blowdown for the vapor or
vapor/liquid relief cases must be considered. Final selection should not be
made prior to consulting with the manufacturer.
3.11.5.6
Pilot Operated Pressure Relief Valves
Pilot operated pressure relief valves should be sized according to
manufacturer’s equations and recommendations. Final selection should not be
made prior to consulting with the manufacturer.
3.11.5.7
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There are no sizing equations published for Section I steam valves. Required
areas can be approximated by using the API steam equations, but this can be
incorrect as manufacturers do not necessarily comply with API effective areas
for Section I (API Standard 526 does not cover Section I safety valves).
Manufacturers will publish equations for a particular safety valve design, which
generally can be used. However, the proper way of specifying Section I safety
valves is to specify the set pressure, required steam capacity and steam
superheat. The manufacturer’s publish tables of their certified capacities on
saturated steam and tables of derating factors for superheat. A safety valve
should be selected from certified tables, not from a sizing equation.
3.11.6
Rupture Disk Sizing
Rupture disks must be sized to comply with Section VIII of the ASME Code and with
local laws and codes. Adequate flow through a rupture disk can be assured if the disk
is sized carefully. Rupture disks should be sized according to manufacturer’s
equations and recommendations. Final selection should not be made prior to
consulting with the manufacturer.
3.11.6.1
Rupture Disks as Primary or Secondary Relief Devices
Section VIII, Division I, of the ASME Boiler and Pressure Vessel Code provides
rules for the use of rupture disk devices for over pressure protection. Within
the scope of the ASME code, rupture disk devices are characterized as
nonreclosing pressure relief devices that may be used to satisfy relief
requirements for over pressure protection of a pressure vessel. Rupture disk
devices must meet the same sizing and set pressure requirements that are
applied to any pressure relief device, such as safety relief valves.
When a rupture disk is installed as a stand-alone pressure relief device, the
size is determined by calculating the discharge area needed to pass the
required rate, using the appropriate equation for the flowing medium. The
rupture disk selected should correspond to the nominal pipe size whose area
is equal to or greater than the required discharge area. In order to confirm that
adequate protection is provided, however, the entire system including inlet and
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discharge piping, and not just the rupture disk, must be checked for the
required flow rate.
There are a number of ways of approaching the problem. From a hydraulic
standpoint, a burst rupture disk is just another resistance in a pipeline. API RP
520, Part I, recommends that a burst rupture disk be taken to have a
resistance coefficient (K=fL/D) of 1.5. If this is done, the capacity of the piping
between the protected vessel and its discharge point can be analyzed using
standard hydraulic techniques, with the disk being a contributor to the overall
system resistance. As in any compressible flow system, the presence of
critical flow restrictions must be taken into account.
A second method is to compute the pressure profile in the inlet piping to the
rupture disk and then compute the rupture disk capacity using the ASME
formula with a 0.62 coefficient of discharge. The piping on the outlet of the
disk must be large enough to result in critical flow across the disk, otherwise
the ASME critical flow formula is not applicable. This method presupposes
that the rupture disk is the flow limiting component in the system.
Selection and sizing of a rupture disk to be used in combination with a relief
valve is governed by the considerations discussed in Section 3.9.3.
3.12
REFERENCES:
1.
Kern, Robert, “Pressure-relief Valves for Process Plants”, Chemical Engineering,
February 28, 1977, pp. 187-194.
2.
Anderson, F.E., “Pressure Relieving Devices”, Chemical Engineering, May 24, 1976, pp.
128-134.
3.
Zahorsky, J.R., “Degradation of Pressure Relief Valve Performance Caused by Inlet
Piping Configuration”, ASME Paper 83-NE-20.
4.
Cox, O.J., and Weirick, M.L., “Sizing Safety Valve Inlet Lines”, Chemical Engineering
Progress, November, 1980, pp. 51-54.
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5.
Politz, F.C., “Poor Relief Valve Piping Design Results in Crude Unit Fire”, 50th Mid-Year
API Refining Meeting, May 13-16, 1985, Kansas City, MO, Volume 64, pp. 192-193.
6.
Nazario, F.N., “Rupture Disks, A Primer”, Chemical Engineering, June 20, 1988, pp. 8696.
7.
“Standards for Direct Contact Barometric and Low Level Condensers”, Fourth Edition,
The Heat Exchange Institute, New York, N.Y.
8.
Lia, Y.S., “Conventional Spring Loaded Safety Relief Valves Subjected to Back
Pressure”, 1989 ASME/JSME Pressure Vessel and Piping Conference.
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TABLE 3.1
API NOZZLE SIZES AND AREAS
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TABLE 3.2
EFFECTIVE AND ACTUAL AREAS/
COEFFICIENTS OF DISCHARGE
API
NOZZLE
API EFFECTIVE
AREA - SQ. CM.
API
Kd
ACTUAL (ASME
NOZZLE AREA - SQ.
CM.
ACTUAL (ASME)
Kd (2)
D
E
F
G
H
J
K
L
M
N
P
Q
R
T
0.71
1.26
1.98
3.25
5.06
8.30
11.86
18.41
23.23
28.00
41.16
71.29
103.23
167.74
.975
.975
.975
.975
.975
.975
.975
.975
.975
.975
.975
.975
.975
.975
0.83
1.47
2.30
3.77
5.89
9.68
13.81
20.52
27.03
32.58
47.87
82.90
120.00
184.64
.851
.851
.851
.851
.851
.851
.851
.851
.851
.851
.851
.851
.851
.851
NOTES
1.
VALUES ARE TYPICAL AND ARE NOT INTENDED TO INDICATE ACTUAL VALUES,
CONSULT THE MANUFACTURER FOR ACTUAL NOZZLE AREA FOR A SPECIFIC
VALVE.
2.
RATED Kd AFTER APPLICATION OF ASME 0.9 DERATING FACTOR. Kds ARE
GENERALLY CONSTANT FOR A GIVEN VALVE DESIGN.
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FIGURE 3.1
CROSS SECTION OF CONVENTIONAL
PRESSURE RELIEF VALVE
1. Base
5. Adjusting
Ring
9. Spindle
13. Nut
17. Adjusting
Screw Nut
21. Bonnet
Gasket
2. Nozzle
6. Adjusting
Ring Pin
10. Spindle
Retainer
14. Spring
18. Eductor
Tube
22. Cap Gasket
3. Disk
7. Disk Holder
11. Bonnet
15. Spring
Washer
19. Cap
23. Guide
Gasket
4. Disk Retainer
8. Guide
12. Stud
16. Adjusting
Screw
29. Adjusting
Ring Pin Gasket
24. Vent Pipe
Plug
Valve figure and description used by permission of Consolidated, a division of Dresser Valve and Controls Division
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FIGURE 3.2-A
OPERATING CHARACTERISTICS OF
CONVENTIONAL SAFETY RELIEF VALVES
IN VAPOR SERVICE
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FIGURE 3.2-B
OPERATING CHARACTERISTICS OF
CONVENTIONAL SPRING OPPOSED PRESSURE RELIEF VALVE
IN LIQUID SERVICE
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FIGURE 3.2-C
OPERATING CHARACTERISTICS OF
LIQUID TRIM PRESURE RELIEF VALVE
IN LIQUID SERVICE
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FIGURE 3.3
CROSS SECTION OF BALANCED BELLOWS
PRESSURE RELIEF VALVE
1. Base
5. Adjusting
Ring
9. Spindle
2. Nozzle
6. Adjusting
Ring Pin
10. Spindle
Retainer
3. Disk
7. Disk Holder
11. Bonnet
4. Bellows
Retainer
8. Guide
12. Stud
13. Nut
17. Adjusting
Screw Nut
21. Bellows
Gasket
18. Disk
Retainer
22. Guide
Gasket
15. Spring
Washer
19. Cap
23. Adjusting
Ring Pin Gasket
16. Adjusting
Screw
20. Bonnet
Gasket
24. Cap Gasket
14. Spring
Valve figure and description used by permission of Consolidated, a division of Dresser Valve and Controls Division
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FIGURE 3.4
CROSS SECTION OF PISTON TYPE
PILOT OPERATED RELIEF VALVE
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FIGURE 3.5
CROSS SECTION OF
DIAPHRAGM TYPE RELIEF VALVE
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FIGURE 3.6
CONVENTIONAL TENSION
LOADED RUPTURE DISKS
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FIGURE 3.7
PRESCORED TENSION
LOADED RUPTURE DISKS
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FIGURE 3.8
COMPOSITE DISKS
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FIGURE 3.9
REVERSE BUCKLING DISK WITH KNIVES
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FIGURE 3.10
PRESCORED REVERSE
BUCKLING DISKS
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FIGURE 3.11
GRAPHITE DISKS
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FIGURE 3.12
RUPTURE DISK
TELLTALE INSTALLATION
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FIGURE 3.13
Kb VERSUS BACK PRESSURE FOR
CONVENTIONAL PRESSURE RELIEF VALVES
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FIGURE 3.14
BACK PRESSURE SIZING FACTOR Kb FOR BALANCED
BELLOWS PRESSURE RELIEF VALVE
(VAPORS AND GASES)
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FIGURE 3.15
TYPICAL BACK PRESSURE CORRECTION FACTOR (Kw)
FOR LIQUID SERVICE BALANCED BELLOWS VALVE
(VAPOR OR LIQUID TRIM)
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FIGURE 3.16
TYPICAL OVERPRESSURE CORRECTION FACTOR (KP)
FOR CONVENTIONAL PRESSURE RELIEF VALVE
IN LIQUID SERVICE
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FIGURE 3.17
RUPTURE DISK BURST PRESSURE AND
MANUFACTURING RANGE TOLERANCES
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4.0.
DETERMINATION OF INDIVIDUAL RELIEF LOADS
4.1.
BASIC PHILOSOPHY
4.1.1
Process Evaluation Basis
Overpressure is caused by an accumulation of material or heat within a process
system. These quantities are accumulating either because they are entering into a
process system at too fast a rate (in excess of design) or the normal means of
removal are operating at too slow a rate (less than design). Where the excess
consists of heat, the expansion or vaporization of liquid or the expansion of vapor
caused by this accumulation of heat must be removed by mass removal. Either
material or heat accumulation then must be removed by protective devices to limit
the system pressure to that safely allowed by the materials of construction of the
vessels, piping and mechanical equipment.
ASME Section VIII code requires that all Section VIII pressure vessels be provided
with protective devices in accordance with ASME Standards. This requirement is
interpreted to mean that each portion of the process systems which can be
independently isolated during operation or shutdown is to be protected by a relief
device. Exceptions are those portions where the isolation valves are under
administrative control and can be assumed to never be in the closed position when
a relief would be required. An example would be a fired heater.
In order to establish the material flows which must be relieved, a baseline for
maximum anticipated process flows must be selected and should be clearly
displayed on the project design documents. One of the three bases listed below is
recommended.
4.1.1.1
Material Balance Rates and Duties
All loads are based on process design material balance rates and duties.
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4.1.1.2
Material Balance Rates and Duties plus Specified Margin(s)
All loads are based on process design material balance rates and duties
plus an allowance for debottlenecking. This allowance is selected as a
percent of material balance rates and is added as a fraction to 1.0 to serve
as a multiplier to the material balance rates. As an example, if 10% of
material balance rates is used as the debottlenecking margin, it would
result in a 1.10 multiplier to material balance rates to establish process
flows to use for determination of the required relief valve rates.
Selection of this basis allows the material balance rates to be increased up
to the specified margin(s) at any time. Keep in mind that the size of
individual relief valves may not be impacted by this margin. Consider the
example where the fire case is significantly larger than any other cause of
relief for an individual pressure relief valve. Within limits, the size of the
margin will not affect the pressure relief valve size. The size of the margin
may or may not affect the size of the relief headers and flare depending on
the overall characteristics of the systems protected.
4.1.1.3
Loads Are Based On Individual Equipment or Process Limitations
o
Pumps
Rate for flow at maximum impeller diameter available for the pump.
Use 130% of rated head at rated speed if the pump curve is not
available. An additional allowance is to be added for turbine drivers
(as discussed below) if they are used to drive the pumps.
o
PSV Inlet/Outlet Sizes
Size for maximum flow based on selected PSV orifice size. Hydraulic
calculations are to follow the design guides in Section 7.
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o
Turbine Drivers
Rate small general purpose turbine drivers at 105% of normal speed
and 105% of design speed for large, special purpose turbines.
Confirm this basis with the vendor when the speed controllers are
selected to assure that sufficient allowances have been taken into
account for the speed controller accuracy and the turbine speed trip
settings.
o
Fired Heaters
Rate fired heaters at 125% of rated firing.
o
Towers
Rate towers for operation up to flood or by the reboiler/condenser
limits, whichever are greater.
o
Heat Exchangers
Rate heat exchangers for clean condition for heat addition and fouled
condition for heat withdrawal.
In all cases where this basis has been selected, where specific design data
for equipment is available on a timely basis, it should be used in place of the
values above to define the maximum limits for equipment. As an example,
where the maximum fuel that can be fired in a heater is limited by the
vendors design, this shall be the maximum heat input.
These bases are included in Chapter 1 of the Flare and Relief System Checklist as
item 3C in Table 1.2.
While selection of the design flow basis with greater design margins increases
operating flexibility, the size and cost of the relief system increase also and need to
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be justified. Selection of the design basis for the relief system must be made at the
beginning of any new project.
4.1.2
Double Jeopardy
Relief events not connected by process, mechanical or electrical causes are not to
be considered simultaneous (double jeopardy). Exceptions to this rule occur when
major events with unacceptable consequences of failure have been identified in
hazops or design reviews. Following review and common agreement, several
simultaneous causes of relief may be added together to form the design basis for a
single relief valve. Examples of applicable systems would be those subject to high
pressures or those containing toxic material.
4.1.3
Utility Losses
Utility failure is often the determining factor for relief header sizing. Two types of
failure may occur - general or partial.
o
General Loss of Utilities
This is the more complicated of the two utility failures types. Evaluation of
the effect of overpressure, attributable to the loss of a particular utility
service, includes the chain of events that could occur and the reaction time
that is involved. The consequences of a major shutdown must be
considered.
o
Partial Loss of Utilities
For a partial utility failure, care must be exercised when taking credit for
equipment unaffected by the failure. In general, credit may be taken for
equipment in continuous parallel service having unrelated energy sources.
No credit may be taken for standby equipment, such as spare pumps or
compressors, even though the energy source may be unrelated. Manual
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cut-in of auxiliaries is operator and time dependent and must be carefully
analyzed before being used as insurance against overpressure.
Following are descriptions of the most common utility system failures considered,
including discussion of the general impact of the failure.
4.1.3.1
Loss of Cooling Water
Complete or partial loss of cooling water supply may be caused by power
failure or equipment breakdown. The layout of the cooling water piping
must also be considered, as it may not be possible to have a complete loss
of cooling water to all units. Types of drivers on cooling water pumps must
be evaluated because the standby pump driver usually has an alternate
energy source.
The maximum quantity relieved under these circumstances is established
by a calculation of heat and material balances around the system.
Usually, loss of overhead cooler/condenser simultaneously brings about
reflux failure. The physical properties of the quantity relieved depend upon
the location and pressure setting of the relieving device on the system.
4.1.3.2
Loss of Electric Power
Loss of electric power can cause a wide range of upsets resulting in
overpressure. Power failure may be local, intermediate or total.
o
Local Power Failure
Local power failure can affect the operation of individual equipment
items, such as air fans, reflux pumps and solenoid valves not on
uninterruptable power supply. These failures are discussed in
further detail under Local Equipment/Operation Failure, Section 4.2.4.
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o
Intermediate Power Failure
Intermediate power failure affects the operation of one electrical
distribution center, one motor control center or one bus. Depending
upon how a group of motors are connected to the power source,
multiple failure conditions may occur. For example, assume that the
reflux pumps of a fractionator are motor driven and connected to the
same bus line. If a bus line failure occurs, the accumulator is
flooded, eventually resulting in a flooded condenser. The preferred
practice is to place motors in complementary service on separate bus
bars to minimize the chance of this type of failure.
o
Total Power Failure
Total power failure is assumed to be 100 percent loss of electric
power except for the Uninterruptable Power Supply (UPS). The
effect of a power failure on all the flows in a unit must be evaluated to
establish the relief flows.
Backup electric power sources must be considered in the analysis of
the effects of electric power failure. For example, critical electronic
instrumentation is generally connected to the Uninterruptable Power
Supply (UPS) to ensure safe and orderly shutdown upon loss of
normal electric power supply.
4.1.3.3
Loss of Steam
Steam or boiler feedwater failure may trigger a series of events that could
result in overpressure. Overpressure may result from the loss of steam to
any turbine driven reflux pump, cooling water pump, or compressor. In
series fractionation where steam is the source of reboiler heat, loss of
steam to the reboiler of one column may cause overpressure in the
following column. Loss of heat results in some of the lower boiling point
material mixing with the bottoms stream and being transferred to the next
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column as feed. In this circumstance, the overhead load of the second
column may consist of its normal vapor load plus the lower boiling point
material from the first column. If the second column does not have
sufficient condensing capacity for the additional vapor load, excessive
pressure could occur.
4.1.3.4
Loss of Fuel
Loss of fuel may cause loss of steam, electric power and/or cooling water,
thus indirectly causing overpressure. The same comments as for loss of
steam to series fractionators apply to loss of fuel in similar service.
4.1.3.5
Loss of Instrument Air
Instrument air failure may be local or general. Depending on the service,
local failure can cause overpressure for various reasons. General failure
causes plant shutdown with each control valve reverting to its fail-safe
position in conformance with safety and overpressure limitation
considerations. The fail-safe characteristics of each control valve are
established as an integral part of overall plant design.
4.1.3.6
Loss of Refrigeration
Loss of refrigeration would be similar to loss of cooling water and may be
the result of power failure, cooling water loss, instrument air failure, control
valve failure or breakdown of refrigeration system mechanical components.
Relieving requirements would be set by a heat and material balance at the
process side relief pressure and temperature, taking into account the
capability of downstream equipment to handle additional vapor loads.
4.1.3.7
Loss of Inert Gas
Loss of inert gas frequently leads to vacuum relief on tankage blanketed
with this media.
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4.1.4
Unsteady State Conditions
Many times the peak relief load for a process occurs during the dynamic transition
from normal operation to some other post-upset condition. Any pressure relief
system evaluation must consider this possibility.
The analysis of transient conditions is difficult and requires considerable judgment.
If an existing plant is under review, input from operating personnel is usually helpful.
The designer must employ a combination of calculations and assumptions in
predicting successive stages of system response after the initiating event.
A dynamic simulation of the system response can be extremely useful, but is usually
difficult to justify in terms of the time and cost involved. Dynamic simulations have
been successfully utilized in generalized case studies, such as exchanger tube
rupture, as an aid to developing the overall approach to a particular type of problem.
Any attempt to model a transient situation with a steady-state simulation technique
is usually of little value and can be misleading.
Some of the relief load calculations which may benefit from dynamic analysis are as
follows:
o
Heat Exchange Tube Rupture
Sudden loss of a tube can be followed by acceleration of the fluid on the low
pressure side and pressurization of that system. The duration and
magnitude of these short term spikes and their impact on the system can
only be assessed by dynamic analysis. This type of analysis is used
frequently for exchanger design pressures in excess of 1000 psig (69 barg,
2
70 kg/cm G).
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o
Fractionation/Distillation Tower Upsets
Steady state analysis of tower relief loads requires conservative assumptions
based on normal tower compositions and temperature profiles. In reality,
tower relief can involve dramatic changes in flows, time lags in heat transfer,
and other key changes which can only be assessed via dynamic analysis.
Unlike tube rupture analysis, which may be required to assure the safety of
the system, tower dynamic analysis is usually completed to demonstrate a
lower relieving requirement. As a result, this type of analysis is done less
frequently, requiring the cost of the analysis to be justified. Assumptions
used in modeling each system must be carefully developed and the
simulation results carefully assessed. Maximum load reductions are
generally obtained with columns having a significant difference in bottom and
overhead product boiling points.
o
Depressuring Impact
Emergency depressuring of a high pressure system such as a reactor loop
can impact the design of both the unit and the flare system. Internal to the
unit, equipment design temperatures may be impacted by the flow of high
temperature gas during depressuring. Flare capacity may also be impacted
by the large, but declining, flow of gas which occurs during depressuring.
Pressurization of the flare header (packing) may help balance the peak flow
quantity resulting from depressuring. Dynamic analysis can be used to
reduce the degree of conservatism which must otherwise be incorporated
into the design of these systems.
4.1.5
Block Valves, Check Valves and Control Valves
Assume that any single check valve will fail in the fully open position, any block
valve will be placed in the worst position, either open or closed (except those under
administrative control), and any control valve will fail or be moved into the position
that maximizes the relief quantity. Assume that inadvertent closing or opening of
only one block valve occurs at any one time.
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4.1.5.1
Block Valves
Blocked discharge situations may apply to compressors, pumps, and other
process systems. Generally, omission of block valves interposed in vessels
in series can simplify pressure relieving requirements. The quantity
relieved may be calculated from the characteristics of the sources of
pressure at set pressure plus overpressure. In the case of heat
exchangers, a blocked outlet may cause thermal expansion or vapor
generation. The effect of frictional pressure drop is taken into account
when determining the relief quantity.
4.1.5.2
Check Valves
A single check valve is usually considered acceptable unless a potential
exists for backflow of high pressure fluid to create pressure that exceed the
test pressures of the equipment. In these cases, consideration should be
given to the provision of a secondary device to minimize the potential for a
reversal of flow. Sizing of pressure reducing (control or relief) facilities on
the suction side of a pump to accommodate the peak flow following failure
of a check valve is not normally recommended since reverse flow may
destroy rotating equipment.
4.1.5.3
Control Valves
Every control valve must be considered as subject to inadvertent operation
from the fully open to fully closed position independent of the failure
position.
4.1.5.4
Control Valve Bypasses
Control valve bypasses are subject to inadvertent opening, although
simultaneous full opening of a control valve and its bypass is normally
considered to be double jeopardy. In some instances, control valve
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bypasses have greater capacity than the control valve itself and may
constitute a controlling factor in over pressuring downstream equipment.
In such cases, installation of a reduced size bypass valve may be a
preferred alternative to overly large pressure relief valves.
If there is a history of having to operate with the bypass open, a correctly
sized control valve should be provided, or the pressure relief system will
have to be sized based upon simultaneous opening of the control valve
and its bypass valve. Also, if the downstream system pressure can exceed
hydrotest pressure (typically 150% of design pressure), and the upstream
2
pressure exceeds about 710 psig (49 barg, 50 kg/cm G), then the relief
system should be designed for simultaneous full opening of the control
valve and bypass valve.
4.1.5.5
Control Valve Limit Stops
In some cases, failure of a control valve to a full-open or full-closed condition
results in an unreasonably sized pressure relief system, or requires that
pressure relief valves be installed in locations where process conditions
make their installation undesirable. In these cases, installation of limit stops
to limit the capacity, or to limit the minimum flow of the control valve may be
considered.
Use of limit stops should be approached carefully. Some basic requirements
relative to limit stops are:
a)
The effect of the proposed limit stops on process operations must be
evaluated. The useful valve capacity after installation of the limit stop
must still meet process requirements, both in terms of maximum and
minimum capacities.
b)
Before limit stops are considered, reduction of valve capacity should
be examined. Often modifying the trim is a more positive method of
reducing valve capacity.
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c)
When they are installed, limit stops should be non-adjustable and
permanently installed. They shall either be installed in the valve trim
or inside the valve actuator. External adjustable stops, calibration of
the positioner or valve bench set or limiting the input signal are not
acceptable methods of limiting valve capacity for pressure relief
system design purposes.
d)
Occasionally, the only viable limit stop designs for a control valve
require external limit stops. In this case, the stops should be tackwelded or otherwise permanently attached so that they cannot be
easily adjusted. Inspection of the control valve for proper limit stop
setting must be included as part of a unit’s start-up procedure and
every time the valve is re-installed after maintenance.
e)
Administrative procedures must be implemented to prevent removal
of the limit stops or modification of the control valve capacity without
first evaluating the effect of the change on pressure relief system
requirements
.
If the absence of limit stops would result in the pressure in a vessel
exceeding 1.5 times the MAWP, limit stops should not be used. In this
case, a smaller control valve or a larger sized pressure relief system should
be provided.
The presence of reduced trim or limit stops and their required settings shall
be shown on unit P&ID’s and a list of valves provided with reduced trim or
limit stops for pressure relief reduction shall be provided to Operations,
Inspection and the refinery Instrument Shop. The valve should be tagged
with a warning sign advising that the reduced trim or limit stops are required
for pressure relief purposes. A statement of why the reduced trim or limit
stops are provided must also be added to the control valve data sheet.
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4.1.6
Control System Response
4.1.6.1
Inlet Control Valves that fail closed
For applications involving inlet control devices that fail in the closed
position, overpressure-relieving devices may not be required; however, the
type of control valve failure must be carefully evaluated.
4.1.6.2
Inlet Control Valves that fail open
For applications involving inlet control devices that may fail in the open
position or may be inadvertently opened as a result of operator action,
pressure-relieving devices may be required to prevent overpressure. The
required relief capacity is the difference between the maximum inlet flow
and the normal outlet flow at relieving conditions, assuming that the other
valves in the system are still in normal operating position (that is, normally
open, normally closed, or throttling). If one or more of the outlet valves are
closed by the same failure that caused the inlet valve to open, the required
relief capacity is the difference between the maximum inlet flow and the
maximum flow from the outlet valves that remain open. All flows should be
calculated at relieving conditions.
In evaluating the effect of control valve failure, consider the following
initiating events and their effects on the controller:
1. Loss of transmission signal to the valve positioner
2. Failure of control valve operating medium (air, hydraulic oil, electricity,
etc.)
3. Process measuring element failure
4. Process measuring element transmitter failure
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4.1.6.3
Effect of other control valves during relief
To avoid subjective decisions involved in evaluating response times and
the effects of controller setting, such as band, reset, and rate, any
automatic control valves that are not under consideration as causing a
relieving requirement and which would tend to relieve the system should be
assumed to remain in the position required for normal processing flow.
Although controllers actuated by variables other than the system pressure
may try to open their valves fully, capacity credit should be taken for such
control valves only to the extent permitted by their normal operating
position, regardless of the valve’s initial conditions. Therefore, unless the
condition of flow through the control valves changes, credit may be taken
for the normal capacity of these valves, corrected to relieving conditions,
provided that the downstream system is capable of handling any increased
flow. Any automatic control valves that are not under consideration as
causing a relieving requirement and which would tend to increase the relief
load should be assumed to function normally.
4.1.6.4
Fail-stationary valves
Some control devices are designed to remain stationary in the last
controlled position when the control signal or operating power fails. Since
predicting the position of the valve at the time of failure is not possible, the
designer should always consider that such devices could be either open or
closed; therefore, no reduction in relief capacity should be considered
when such devices are used.
4.1.7
Operator Intervention
The decision to take credit for operator response in determining maximum relieving
conditions requires consideration of the complexity of the process controls,
manpower available to respond to the emergencies, and training of personnel. The
required response should be unambiguous and taken as a result of clear indications
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that this response is required. The minimum response time to be taken is 10
minutes. In order to take this credit the following three conditions must be met:
o
No relief occurs for 10 minutes after the upset becomes evident to the
operator
o
Evidence of the upset normally comes from an alarm unaffected by the
initiating event.
o
The operator must be able to correct or stop the event before relief with a
simple response from one location.
The effectiveness of this response will need to be considered as well. A correct
response may reduce but not eliminate the problem completely.
4.1.8
Heat Transfer Equipment Performance
4.1.8.1
Air Cooled Exchangers
The heat removal capacity of an air cooled heat exchanger is not totally
lost when the fan is lost due to mechanical problems or power failure. A
credit of 25% of the normal design duty may be taken because of natural
convection heat transfer. This allowance may be larger but no further
credit should be taken unless confirmed by calculation or by the equipment
manufacturer. Where the loss of cooling is caused by failure of louvers to
a closed position or manual closure of louvers no credit should be taken.
4.1.8.2
Shell and Tube Exchangers
Shell and tube exchangers may be thermally rerated under upset
conditions using the normal heat transfer coefficient, unless otherwise
indicated in this manual. Rerating of reboiler for reduced duty resulting
from a decrease in MTD should follow the guidelines provided in Section
4.3.3.2 or be rigorously calculated.
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4.1.8.3
Fired Heaters
In general, a fired heater is not susceptible to a temperature pinch, and will
be expected to operate at its design duty under relief conditions. No credit
is taken for a temperature controller acting to reduce firing, due to the
potential lag in response time and the possibility that the controller could be
in the manual mode.
Under any circumstance in which process circulation through the heater is
lost, the duty is assumed to go to zero for the purpose of relief load
calculations and methods must be in place to prevent heater tube failure.
4.1.9
Use of DIERS Methodology
In recent years considerable research into the behavior of pressure relief systems in
two-phase, reacting or flashing services has been done by the AIChE’s Design
Institute for Emergency Relief Systems (DIERS). This group consists of members
from sponsoring Chemical Companies and has primarily been charged with
research activities on pressure relief of reacting systems. Several papers on this
group’s activities have been published (1, 2, 3, 4) and the results from the DIERS
Program are summarized in the book titled “Emergency Relief System Design Using
DIERS Technology” (11). At the completion of the DIERS project, the DIERS Users
Group was formed to continue the development of improved emergency relief
system design methods and to exchange information between member
organizations.
DIERS technology includes bases for determining when multi-phase flow in the
emergency relief system must be considered, how to size the emergency relief
device and associated piping when multi-phase flow must be considered, and how
much and at what rate condensed phase will be carried over into the emergency
relief system so that downstream equipment such as K.O. drums and quench drums
can be sized. This work has not yet attained a status of general use in the refining
industry. In its current form, DIERS technology is very complex and depends greatly
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upon pilot plant or test bench studies to properly characterize a fluid in terms of the
DIERS correlations. While the need for this technology has been demonstrated in
the chemical industry where complex exothermic reactions occur, there has not yet
been demonstrated a need for it in the refining industry where such reactions are
rare.
While use of DIERS technology does not now appear to be required, or practical,
further developments may make routine use of it in sizing pressure relief systems for
flashing or boiling liquids possible. In the meantime, API RP 520, Part I still
recommends use of conventional two-phase sizing techniques, but advises of the
existence of DIERS. Several refiners have begun to use DIERS or similar
methodologies (see Section 3.11.5.4) when sizing relief valves for two-phase flow.
4.2.
CAUSES OF OVERPRESSURE
4.2.1
General
Table 4.2 lists 16 conditions that have been identified in API RP 521 (Third Edition),
Table B-1, as causing need for protective devices as well as the bases to quantify
the required relief rates. A detailed discussion of these items is available in Chapter
3 of API RP 521. The discussion that follows has focused on the major refinery
elements such as fractionation systems, reactor loops, blocked discharges in the
process, mechanical equipment, heat exchangers, plant fire, chemical reactions and
atmospheric storage tank protection.
An excellent start to establishing the required number, location and approximate
size of relief valves in a process unit is to look at the Piping and Instrument Diagram
(P&ID) for an identical or similar process unit. However, the final selection of relief
valves may be slightly different because of potential differences in design
philosophy.
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4.2.2
Operator Error
Operator error is a potential factor causing plant upsets and overpressure. The
most common error is the inadvertent closing or opening of block valves. Proper
training, safe plant practices, posting of legible instructions and warning signs, and
provisions for valves locked or sealed in the open or closed position (administrative
control) are among the preventative measures which minimize this type of
contingency.
4.2.3
Utility Failure
The consequences that may develop from the loss of any utility service, whether
plant wide or local, must be carefully evaluated. The normal utility services that
could fail and a partial listing of affected equipment that could cause overpressure
are given below:
o
Electric
- Pumps for circulating cooling water, boiler feed, quench, or reflux.
- Fans for air-cooled exchangers, cooling towers, or combustion air
- Compressors for process vapor, instrument air, vacuum, or refrigeration.
- Instrumentation
- Motor-driven valves
o
Cooling Water
- Condensers for process or utility service
o
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-
Coolers for process fluids, lubricating oil, or seal oil
-
Jackets on rotating or reciprocating equipment
Instrument Air
-
Transmitters and controllers
-
Process-regulating valves
Alarm and shutdown systems
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o
4.2.4
Steam
- Turbine drivers for pumps, compressors, blowers, combustion air fans, or
-
electric generators.
Reboilers
-
Reciprocating pumps
Equipment that uses direct steam injection
Eductors
o
Fuel Oil/Fuel Gas
- Boilers
- Reheaters (reboilers)
- Engine drivers for pumps or electric generators
- Compressors
- Gas turbines
o
Inert Gas
- Seals
- Catalytic reactors
- Purge for instruments and equipment
Local Equipment/Operation Failure
4.2.4.1
Reflux Failure
Reflux failure is one of the more common causes for overpressure. The
loss of reflux usually results in overhead condenser flooding. Relief
devices are sized to handle relief loads based on total loss of cooling, if a
condenser can flood before operator response to eliminate the upset
condition is likely to occur. Therefore, the accumulator shall be checked
for the flooding potential during this contingency.
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4.2.4.2
Abnormal Heat Input - Reboiler
Abnormal heat input to a reboiler may cause overpressure when vapor
generation capacity exceeds condensing capacity. Overpressure can
occur when reboilers are new, in clean service, and conservatively sized or
when temperature control fails.
4.2.4.3
Heat Exchanger Tube Failure
Heat exchanger tubes are subject to failure from a number of causes. If
the hydrostatic test pressure of the low pressure side cannot be exceeded,
relief protection is not required. Since typically hydrotest pressures are
equal to 150 % of design pressure, a relief valve is not required if the low
pressure side (system) design pressure is greater than or equal to twothirds of the operating pressure of the high pressure side. Installation of a
pressure relief valve is not warranted if the piping and downstream
equipment on the low pressure side can handle material from the high
pressure source without exceeding 110 percent of their design pressure.
Where overpressure protection is required, the tube rupture flow rate into
the low pressure side should be based on an orifice coefficient of 0.6 and
an area equal to twice the cross sectional area of one tube, using the
operating upstream pressure and 110 percent of the downstream design
pressure. If this calculated flow exceeds the total normal flow on the high
pressure side, the rate at which the high pressure side systems pressure
will decline should be considered in calculating an appropriate flow rate
though the tube rupture. Allowance must be made for any liquid which
could flash.
4.2.4.4
Condenser Failure
Air condenser failure may cause overpressure in upstream equipment. For
air cooled condensers the relieving quantity may be reduced due to cooling
by natural convection. A value of 25 percent reduction should be assumed
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unless the equipment vendor verifies a larger percentage. If the piping
system is unusually large and bare, the effect of heat loss from the piping
to the surroundings can also be considered.
The relief requirement for partial condenser failure is the difference
between the total incoming rate and the outgoing liquid rate at relieving
conditions.
4.2.4.5 Louver Failure
If a louver is inadvertently closed, no credit may be taken for natural
convection and the total incoming vapor must be vented. Closure may
result from automatic control failure, from destructive vibration of
mechanical components or from operator error. The relief requirements for
total condenser failure are the total incoming vapor calculated at relieving
conditions with the prevailing heat input.
4.2.4.6
Loss of Absorbent Flow
Loss of absorbent flow to a gas absorber, where a major component of the
inlet vapor is removed, could result in overpressure. This results if the
downstream equipment does not have the capability to pass the increased
flow. When sizing the relief device, credit is taken for the design capability
of downstream equipment.
4.2.4.7 Automatic Process Control Failure
Automatic process control failure requires an analysis of the process
variables being controlled, which are pressure, temperature, flow and liquid
level. By definition, the control valves are located at the inlet or discharge
of a process system. The misoperation of any one control valve could
result in overpressure due to inadvertent valve opening or blocked
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discharge conditions. General control valve failure reverts to instrument air
type failure.
4.2.5
External Fire
Fire provides an unanticipated energy input to a system, which results in
overpressure by thermal expansion or vaporization of the retained fluid. Section 4.8
discusses methods for evaluating relief loads due to external fire relief.
4.2.6
Depressurization
4.2.6.1
Fire
Depressurization is used to reduce the pressure of vessels, equipment and
piping below normal operating pressure to prevent rupture caused by
localized fire heating.
See API RP 521, 3rd Edition, 1990, paragraph 3.19 for further discussion.
4.2.6.2
Chemical Reactions
Exothermic reactions accelerate with rising temperature producing
extremely high rates of energy release. Large volumes of
noncondensibles may be produced if the temperature rises to excessive
levels and decomposition reactions begin to occur. Rapid relief is a
requirement in coping with possible runaway reactions.
A possible method to prevent destructive overpressures in such reactions
is to add enough volatile fluid to absorb excess reaction heat and cool the
reacting material to safe levels. If sufficient cooling results from the
depressuring, vapor depressuring is an even better means to avert
runaway reactions.
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4.2.7
Thermal Expansion
Thermal expansion or hydraulic expansion is the increase in liquid volume caused
by an increase or decrease in temperature. It can result from several causes, the
most common of which are the following:
o
Piping or vessels are blocked in while they are filled with cold liquid and are
subsequently heated by heat tracing, ambient heat gain, or fire.
o
An exchanger is blocked in on the cold side with flow on the hot side.
o
Piping or vessels are blocked in while they are filled with liquid at nearambient temperature and are heated by direct solar radiation.
In certain installations, such as cooling circuits, the processing scheme, equipment
arrangements and methods, and operating procedures make feasible the
elimination of the hydraulic expansion relieving device, which might normally be
required on the cooler fluid side of a shell-and-tube exchanger. Typical of such
conditions would be multiple-shell units with at least one cold fluid block valve locked
open on each shell, and a single-shell unit in a given service where the shell can
reasonably be expected to remain in service, except on shutdown. In this instance,
the cold-fluid block valves on the exchanger unit should be posted with signs
stipulating the proper venting and draining procedures when shutting down and
blocking in. Such cases are acceptable and do not compromise the safety of
personnel or equipment, but the designer is cautioned to review each case carefully
before deciding that a relieving device based on hydraulic expansion is not
warranted.
Liquid expansion rates for the sizing of relief devices that protect heat exchangers,
condensers, and coolers against thermal expansion of trapped liquids can be
approximated using the following formula:
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Q = 0.00928 (BH) / (GC)
4.1 (English)
Q = 0.0167 (BH) / (GC)
4.1 (Metric)
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Where:
Q
=
B
=
3
3
flow rate at the flowing temperature, ft /hr (m /hr)
Cubical expansion coefficient per degree Centigrade for the liquid at
the expected temperature. This information is best obtained from the
process design data; however, Table 4.1 shows typical values for
hydrocarbon liquids and water at 60 °F (16 °C).
H
=
Total heat transfer rate, Btu/hr (kJ/hr, kcal/hr). For heat exchangers,
this can be taken as the maximum exchanger duty during operation.
G
=
Density at 60 °F (16 °C), lb/ft3 (kg/m3 )
C
=
Specific heat of the trapped fluid, Btu/hr °F (kJ/kg °C, kcal/kg °C)
TABLE 4.1
CUBICAL EXPANSION COEFFICIENT
TYPICAL VALUES FOR HYDROCARBON LIQUIDS
AND WATER AT 16 °C
3-34.9 deg API gravity
35-50.9 deg API gravity
51-63.9 deg API gravity
64-78.9 deg API gravity
79-88.9 deg API gravity
89-93.9 deg API gravity
94-100 deg API gravity and lighter
Water
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0.00072
0.0009
0.00108
0.00126
0.00144
0.00153
0.00162
0.0018
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4.2.8
Chemical Reaction
The methodology for determining the appropriate size of an emergency vent system
for chemical reactions was established by DIERS (Design Institute for Emergency
Relief Systems).
The DIERS methodology is based on:
o
Defining the design basis upset conditions for the reaction system.
o
Characterizing the systems through bench scale tests simulating the design
basis upset conditions.
o
Using vent sizing formula which account for two phase gas/liquid vent flow.
The design basis upset conditions are process specific, but generally include one or
more of the following:
o
External fire
o
Loss of mixing
o
Mischarge of reagents
Reaction rates are rarely known, therefore, bench scale tests simulating the design
basis upset condition are usually required. Test equipment is available for this
purpose. With the information obtained from the bench scale tests, the system can
be characterized by one of the following terms:
a)
Tempered
Tempered systems are those in which the unwanted reaction produces
condensable products and whose rate of temperature rise is tempered by
liquid boiling at system pressure. Typically, tempered systems are liquid
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phase reactions in which a reactant (or solvent) is a major portion of the
reactor contents.
b)
Gassy
Gassy systems are those in which the unwanted reaction produces noncondensable products and whose rate of temperature rise is not tempered by
boiling liquid. Gassy systems can be either liquid phase decompositions or
vapor phase reactions.
c)
Hybrid
Hybrid systems are those whose rate of temperature rise due to an
unwanted reaction, may be tempered by liquid boiling at system pressure,
but can also give rise to the generation of non-condensable gas.
Following characterization of the system, the appropriate vent sizing formula can be
selected. The reader should be cautioned that this is an area with rapidly changing
technology and the most current technology should be used, if available.
If the bench scale simulations indicate the potential for an explosion, the
considerations for an explosion or detonation should be considered.
Where feasible, a pressure relief device should be used to control overpressure.
Where this is infeasible, other design strategies may be employed to control
equipment over-stressing. These strategies may include using safety systems such
as: automatic shutdown systems, other forms of reactions that generate heat
(dilution of strong acids) should also be evaluated, inhibitor injection, quench, deinventorying, alternative power supplies, and depressuring. When this approach is
taken, the reliability of the protective system(s) should be addressed in a formal risk
analysis.
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4.2.9
Miscellaneous
4.2.9.1
Entrance of Volatile Material Into the System
o
Water into Hot Oil
Although the entrance of water into hot oil remains a source of
potential overpressure, no generally recognized methods for
calculating the relieving requirements are available. In a limited
sense, if the quantity of water present and the heat available in the
process stream are known, the size of the relief valve can be
calculated like that of a steam valve. Unfortunately, the quantity of
water is almost never known, even within broad limits. Also, since
the expansion in volume from liquid to vapor is so great
(approximately 1:1400 at atmospheric pressure) and the speed of
vapor generation is essentially instantaneous, it is questionable
whether the valve could open fast enough to be of value. Normally,
no pressure-relieving device is provided for this contingency. Proper
design and operation of the process system are essential in attempts
to eliminate this possibility. Avoiding water-collecting pockets and
installing proper steam condensate traps and double blocks and
bleeds on water connections to hot process lines are some
precautions that can be taken.
o
Light Hydrocarbons into Hot Oil
The information above for water into hot oil applies to the entrance of
light hydrocarbons into hot oil even though the ratio of liquid to vapor
volume may be considerably less than 1:1400.
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Page 1 of 3
TABLE 4.2
BASES FOR RELIEF CAPACITIES UNDER SELECTED CONDITIONS
ITEM
NO
PRESSURE RELIEF DEVICE
CONDITION
LIQUID RELIEF
1
Closed outlets on vessels
2
Cooling water failure to
condenser
Total vapor to condenser at relieving
conditions
3
Top-tower reflux failure
Total incoming steam and vapor plus
that generated therein at relieving
conditions less vapor condensed by
side stream reflux
4
Side stream reflux failure
Difference between vapor entering
and leaving section at relieving
conditions
5
Lean-oil failure to absorber
None, normally
6
Accumulation of noncondensables
Same effect in towers as found for
Item 2; in other vessels, same effect
as found for Item 1
7
Entrance of highly volatile
material
For Towers - Usually not predictable
Consider:
o Water into hot oil
Maximum liquid;
pump-in rate
VAPOR RELIEFa
Total incoming stream and vapor plus
that generated therein at relieving
conditions
For Heat Exchangers - Assume an
area twice the internal cross-sectional
area of one tube to provide for the
vapor generated by the entrance of
the volatile fluid due to tube rupture.
o Light HC into hot oil
a
Consideration may be given to the suppression of vapor production as the result of the device’s relieving
pressure being above operating pressure, assuming constant heat input. (Procedures for sizing pressure relief
devices are represented in API Recommended Practice 520, Part I).
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Page 2 of 3
TABLE 4.2 (Continued)
BASES FOR RELIEF CAPACITIES UNDER SELECTED CONDITIONS
ITEM
NO
CONDITION
PRESSURE RELIEF DEVICE
LIQUID RELIEF
8
9
Overfilling storage or surge
vessel
Failure of automatic controls
10
Abnormal heat or vapor input
11
Split exchanger tube
12
Internal explosions
13
Chemical reaction
14
Hydraulic expansion
o Cold fluid shut in
o Lines outside process
area shut in
VAPOR RELIEFa
Maximum liquid;
pump-in rate
Must be analyzed on a case by case
basis
Estimated maximum vapor generation
including non-condensables from
overheating
Steam or vapor entering from twice
the cross-sectional area of one tube;
also same effects found in Item 7 for
exchangers
Not controlled by conventional relief
devices but by avoidance of
circumstances
Estimated vapor generation from both
normal and uncontrolled conditions
See API RP 520,
Fifth Edition,
Appendix C.2
See API RP 520,
Fifth Edition,
Appendix C.2
a
Consideration may be given to the suppression of vapor production as the result of the device’s
relieving pressure being above operating pressure, assuming constant heat input. (Procedures for
sizing pressure relief devices are represented in API Recommended Practice 520, Part I).
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Page 3 of 3
TABLE 4.2 (Continued)
BASES FOR RELIEF CAPACITIES UNDER SELECTED CONDITIONS
ITEM
NO
CONDITION
PRESSURE RELIEF DEVICE
LIQUID RELIEF
15
Exterior fire
16
Power Failure (Steam,
electric, or other)
VAPOR RELIEFa
Estimate by the method given in
API RP 520, Appendix D.5
o Fractionators
All pumps could be down, with
the result that reflux and cooling
water could fail
o Reactors
Consider failure of agitation or
stirring, quench or retarding
stream; size valves for vapor
generation from a runaway
reaction
o Air-cooled exchangers
Fans could fail, size valves for the
difference between normal and
emergency duty
o Surge vessels
a
Maximum liquid;
inlet rate
Consideration may be given to the suppression of vapor production as the result of the device’s relieving
pressure being above operating pressure, assuming constant heat input. (Procedures for sizing pressure
relief devices are represented in API Recommended Practice 520, Part I).
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4.3
FRACTIONATION AND DISTILLATION
4.3.1
System Description
Columns are usually protected by a pressure relief valve (or valves) mounted on the
top head or overhead line. This means that a single pressure relief installation is
used to protect the tower itself plus all associated equipment. This can include the
reflux drum, side stream strippers, overhead condenser, reboiler, and possibly the
feed and product exchangers, depending on the piping configuration and location of
the control valves. A supplementary pressure relief valve may be provided on the
reflux drum but usually only for fire purposes.
4.3.1.1
Heat and Material Balance Envelope
Establish the H&M balance envelope for the tower system to be evaluated
including (see attached examples in Appendix B-1):
•
Tower overhead condenser, accumulator and reflux system.
•
Tower bottom reboiler, exchanger or fired heater.
•
Pump around cooling system.
•
Preheater, if a fired heater is located at the tower inlet.
•
Side-draw strippers and surge pots.
All streams leaving the envelope shall be saturated liquid or vapor. Do not
include additional heat recovery exchangers in the tower feed system. The
incremental heat to the feed will be separately reviewed and incorporated in
the feed enthalpy (see Section 4.3.3.2.). Also do not include product coolers.
Assume the overhead product from the reflux pump is a saturated liquid from
the reflux drum.
4.3.1.2
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Normal Heat and Material Balance
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4.3.2
a)
Establish a normal heat and material balance for tower in and out of
the envelope.
b)
API and UOP K values are required for calculating enthalpy from charts
if they are not available from a simulation.
Causes of Overpressure
Some potential causes for overpressure leading to the need for relief are listed below,
in three main categories:
a)
Operational Failures
-
b)
Loss of Cold Feed
Reflux Failure
Pumparound failure
Loss of Quench
Utility Failure (power, steam, cooling water, instrument air)
Control Failure
Abnormal Heat Input
Absorbent Flow Failure
Blocked Outlet
Compositional Changes
- Accumulation of Non-Condensables
- Loss of Heat in Series Fractionation System
c)
Other Conditions
- Fire
- Exchanger Tube Rupture
Cases listed as operational failures generally present the most severe upsets to
column operating, and will frequently govern pressure relief valve sizing.
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4.3.3
Heat and Material Balance Considerations for Upset Conditions
4.3.3.1
Basic Assumptions for Relief Case Heat and Material Balance
Develop a heat (enthalpy) balance in and out the tower envelope to calculate
the relief load. The tower relief load is calculated according to the following
assumptions and basis.
a)
Detailed hydraulic calculations will not normally be performed to
determine a reduced feed rate. Instead, assume all feeds are entering
the tower envelope at the same mass rates as normal condition, unless
feed pumps stop or maximum source pressure (e.g., pump shutoff
pressure) is not clearly higher than the relieving pressure.
b)
If the feed flow condition is clearly in the mid-range between normal
and zero flow, and the difference in the calculated relief flows for these
two alternative conditions is large enough to be critical, a hydraulic
calculation can be considered to estimate the actual feed rate
expected.
c)
When all feeds are assumed the same as normal, assume all products
leave the tower envelope at the same mass rates as normal. Assume
all product compositions are the same as normal.
Even if the product pumps stop, the products are assumed to leave the
tower envelope, accumulating at the individual product surge vessels
such as tower bottom, side stripper and overhead accumulator. If side
product surge capacity is low, it may impact reboiler pinch credit which
is described in paragraph b) of the Reboiler Duty section.
d)
When a zero feed condition is assumed, assume no product out. Do
not assume any liquid level pulldown.
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When a reduced feed rate is assumed, due to stoppage of one of the
feeds, reduce the product flow rates to make a mass balance by
approximately the same quantity as that "product" is contained in the
feed.
e)
Assume unlimited liquid is available at the top tray composition to
generate relief vapor out of excess heat. Composition of this liquid is
the same as the normal reflux composition. The accumulation rate is
calculated by assuming vaporization by unbalanced heat.
4.3.3.2
Heat Balance for Upset Conditions
Tower heat balance for Upset Conditions is calculated by use of the following
assumptions:
o
Stream Pressure
Establish the pressure profile of the fractionating tower system during
relief. Assume the pressure at the pressure relief valve location is
equal to the Maximum Allowable Accumulated Pressure of the vessel.
Assume the fractionating tray pressure drop is the same as during
normal operations.
o
Stream Compositions and MW
The compositions and molecular weights of the tower tray liquid
streams at various locations in the fractionating tower system are
assumed to be the same during relief conditions as during normal
operations.
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o
Liquid Product Temperature and Enthalpy
Calculate liquid temperature and enthalpy for each product (except OH
liquid) as a saturated liquid at the relieving pressure at each product
location.
o
OH Product Temperature and Enthalpy
Assume the OH product exits the tower at the relieving temperature
(see Accumulation Temperature and Enthalpy paragraph below).
Assume the OH liquid product leave as a saturated vapor. If steam is
present in the normal overhead vapor, assume this steam also exits
the tower envelope at the relieving temperature.
o
Feed Temperature and Enthalpy
The feed temperature is only calculated for reference from the feed
enthalpy calculation. Assume same as normal feed temperature if the
feed enthalpy is not changed. Feed enthalpy is calculated by adding
increase (or decrease) of heat duty during relief condition in the feed
preheating exchangers (see Other Exchanger Heat Duty paragraph
below).
o
Accumulation Temperature and Enthalpy
The accumulation vapor and liquid enthalpy are calculated at the reflux
stream dew point. Relieving temperature is considered as the 100%
vapor temperature, which could be different from the top tray vapor
temperature due to presence of non-condensables and steam.
o
Reboiler Duty
Assume the reboiler duty during relief is same as the normal operating
duty unless the following conditions apply during relief:
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a)
If flow to reboiler stops, reboiler duty is zero. Refer to Heater
Duty Section for fired reboilers.
b)
If the clean surface heat transfer with reduced LMTD is less
than the normal duty, consider a reboiler pinch. Pinch
calculations will only be considered when the calculated load
with normal duty is significant. No pinch credit is taken for fired
reboilers. Caution should be used in calculating a reboiler pinch
to assure that bottoms composition changes resulting from
lighter components flowing down the column as a result of
reduced reboiler duty are considered in the calculation. This
potential reboiler composition impact can be evaluated
conservatively as follows for a simple column (one feed,
overhead product only).
Loss of Feed and Reflux: Use the normal reboiler composition
to evaluate the pinch.
Loss of Reflux: Use the normal reboiler composition to initially
evaluate the pinch. If the calculated relief load is less than or
equal to the net normal overhead product, then the reboiler duty
can be recalculated based on the feed composition or the relief
load set at the normal net overhead product rate. If
recalculated, the lesser of the recalculated load or the net
normal overhead product rate should then be used as the relief
load.
Loss of Overhead Condenser (Partial or Complete, w/o Loss of
Reflux): Use the normal reboiler composition to initially
evaluate the pinch. If the calculated relief load is less than or
equal to the normal total column overhead vapor rate, then the
reboiler duty should be recalculated based on a tray
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composition calculated by the maximum volume of liquid
displacement during the relief case to the reboiler, as follows:
i)
Estimate the net tray liquid displacement volume during
the assumed 10 minute relieving period as: the total
reflux returned + normal product rate (based on
continuation of feed contribution) - calculated relief load.
Limitations on the availability or loss of reflux during this
10 minute period should be taken into account in
calculating the total volume of reflux returned to the
column. However, the other two rates are assumed to
be constant during this period.
ii)
Divide the volume of liquid from i) by the estimated
volume of liquid per tray based on a 2 in (50 mm)
average liquid height per tray over the total column
diameter. This conservatively calculates the number of
tray volumes of liquid that could be displaced into the
bottoms.
iii)
Recalculate the relief load using the normal composition
for a tray located at or above the number of trays from
the bottom calculated in ii). As a simplification, use of
the feed composition for trays below the feed level or
the reflux composition for trays above the feed can be
considered.
Due to the added effort required to complete this trial and error
calculation, reboiler pinch calculations will not normally be done
for cases involving loss of overhead condenser duty without
loss of reflux. Exceptional conditions which may warrant the
calculation include columns with very large relief loads and
columns with a large difference between operating and relief
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pressure, such that process side temperature changes are
likely to be substantial.
Loss of Feed - This pinch calculation is very similar to the
calculation for the loss of overhead condenser duty, except the
volume calculated in step i) is reduced by the normal product
rate (due to feed loss).
Calculations for more complex columns should be based on
similar logic.
o
c)
In cases where as-built system data is available and detailed
calculations are justified, hydraulic calculations can also be
used to establish any supply limitations of heat transfer media,
as a supplement to the pinch calculations in step b.
d)
To complete reboiler pinch calculations for steam reboilers, use
the saturated steam temperature at the normal steam pressure
(from utility design information).
Overhead Condenser Duty
a)
As a starting point, assume (conservatively) that condenser
duty during relief is same as the normal duty unless stoppage
(or reduction) of coolant flow (air or water) occurs during relief.
b)
It is permissible to recalculate the condenser duty to account for
increased heat transfer at the elevated process side
temperature and pressure, if the relief load based on normal
condenser duty is judged excessive or is over the existing relief
valve and/or header capacity. However, this calculation is
typically not required, because the controlling relief load most
often involves loss of overhead condensing capacity.
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c)
Assume natural convection cooling of air condenser at total fan
failure is 25% of normal condensing duty unless a more
accurate number is obtained from the manufacturer. If an
automatic controlled louver is provided for temperature control
and the louver fails closed during the contingency, no natural
convection credit shall be taken.
o
d)
Water condenser duty during cooling water failure shall be
adjusted for any reduction in cooling water flow as defined in
the utility failure design basis.
e)
For reflux failure, assume condenser duty is zero as the result
of liquid backup unless the accumulator has 10 minutes surge
time above the independent high level alarm before the
condenser is flooded (unless the operator is alerted by a
previous system alarm). In the event of a power failure,
calculated surge time should be based on the normal liquid
level.
f)
For relief cases where there is a loss of non-condensables
offgas flow from the reflux drum assume the loss of all
condenser duty as the result of the accumulation of noncondensables gas in the reflux drum and condenser.
Other Exchanger Heat Duty
Heat transfer duties during relief should be determined for heat transfer
equipment in the protected system as follows:
a)
Use zero heat transfer duty when there is total loss of flow
through the exchangers.
b)
Normal duty can be used if any change tends to reduce relief
load. However, if the duty during relief tends to increase relief
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load or there is a need to reduce the calculated relief load, then
the exchanger duties can be recalculated for relief conditions.
Calculate the heat transfer duty during relief by prorating the
normal duty by a factor equal to the ratio of LMTD at relief
conditions and LMTD at normal operation, except for unusual
conditions, such as phase changes, where it may be necessary
to rerate the exchanger due to a significant change in heat
transfer coefficients.
c)
o
If no information is available from the manufacturer, assume 25
percent of the normal duty is obtained by natural convection if
aircooler fans are lost on power failure.
Fired Heater Duty
a)
If the process stream flows to the heater, and firing is not
interrupted, use 100 percent of the normal heat transfer duty,
except for the fuel control valve wide-open case. If vendor
information is available, use the maximum firing capacity
provided by the vendor. Otherwise, use 125 percent of heater
design duty for the relief case when the fuel control valve fails
wide open.
b)
If the fuel shutoff valve is closed (or if air flow is lost) but
process flow continues, assume the following percentage of
duty reduction, unless a more accurate number is obtained
from manufacturer. (Note: Fuel shutoff safety systems must be
in place to prevent a hazardous condition if air flow is lost).
•
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Heat input is reduced to 25 percent of normal duty for fired
heaters with low density sidewall insulation, such as
ceramic fiber.
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•
Heat input is reduced to 50 percent of normal duty for fired
heaters with high density sidewall insulation, such as fire
brick or castable refractory.
c)
Use zero heat transfer duty, if there is total loss of flow
through the fired heater and natural convection flow is not
possible.
While boilup continues temporarily at a reduced rate, a
significant portion (sometimes all) of the retained liquid in the
tubes is required to heat up and pressure the column. Mass
flow rate from the heater will be reduced and fuel shutoff is
likely to occur (if not, tube rupture is of greater concern than
any relief load generated). Some overhead air-cooling is also
retained. This cooling has a greater impact due to the reduced
rate of vapor generation.
4.3.4
Maximum Capacity
For a distillation column, the considerations related to maximum capacity of the
tower should be in accordance with the decisions taken in regard to paragraph 4.1.1
If the maximum capacity is higher that the basis selected for the heat and material
balances, the usual approach is simply to factor up the heat and material balances
for the tower and then compute the governing relief case.
In the case of a multi-product fractionator, the definition of ultimate capacity is not a
clear-cut matter. Column flooding may provide some indication, but a column with
one or more side draws and/or intermediate reflux circuits offers considerable
flexibility in redistributing column traffic to avoid bottlenecks. The same argument
may be applied to perceived limitations in the overhead circuit, which can be
unloaded by changing product draw rates or increasing pump around cooling duties.
For these reasons, the maximum throughput is most likely to be determined by
limitations on the charge heater, or on the upstream reactor or coke drum, etc.,
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which feed the fractionator. These types of limitations should have already been
identified when defining the unit evaluation basis.
4.3.5
Determination of Relief Loads
Relief loads for fractionation and distillation columns are calculated using the
procedures in this section. A sample calculation is provided in Appendix B-1 and the
spreadsheet program for this calculation is provided in Appendix D.
4.3.5.1
Accumulation
Accumulation is calculated by using the assumptions described in paragraph
4.3.3.1. It is assumed that the unbalanced heat generated during relief
condition vaporizes top tray liquid.
ΣW F = ΣW P
WA
4.2
= (ΣW FHF + ΣQI) - (ΣW PHP + ΣQO)
4.3
LA
Where:
Chap4-r1.doc
WA
=
Accumulation rate, lb/h (kg/h)
ΣW F
=
Summation of rates of all feed streams, lb/h (kg/h)
ΣW P
=
Summation of rates of all product streams, lb/h (kg/h)
LA
=
Latent heat of vaporization of the accumulation, Btu/lb
(kJ/kg, kcal/kg)
ΣQI
=
Summation of all heat inputs, Btu/h (kJ/h, kcal/h)
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ΣQO
=
Summation of all heat removed, Btu/h (kJ/h, kcal/h)
ΣW FHF
=
Summation of the products of each feed rate times its
enthalpy, Btu/h (kJ/h, kcal/h)
ΣW PHP =
Summation of the products of each product rate times
its enthalpy, Btu/h (kJ/h, kcal/h)
4.3.5.2
Relief Rate
The relief load (on a mass basis) for the fractionation tower is the summation
of the following streams:
o
Accumulation calculated in paragraph 4.3.5.1.
o
Overhead liquid product calculated as saturated vapor. This stream is
assumed to be vaporized at relief pressure from the normal liquid OH
product. If the sum of accumulation and overhead product liquid rates
are negative, then use zero as the sum. A check must then be made
to determine if off gas or steam must be relieved.
o
Off gas (or normal vapor product) if vapor flow is blocked. When vapor
flow is assumed to continue , do not add this stream.
o
Steam, if it is assumed that steam is not condensed and leaves the
tower envelope (refer to Section 4.3.3.2.) and other non-condensables
which may be introduced abnormally into the tower due to a process
upset.
4.3.5.3
Relief Gas Physical Properties
Average physical properties for relief gas are calculated based on relief
streams defined in Section 4.3.5.2. The physical properties required for relief
and flare work are:
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- Molecular Weight (MW)
- Temperature, °F (°C)
- Specific heat ratio (k = Cp/Cv), no units
- Gas compressibility factor (Z), no units
- Gas viscosity, centipoise
- Low Heating Value (LHV), Btu/lb (kJ/kg, kcal/kg)
4.3.5.4
Impact of Upset Conditions on Relief Requirements
o
Loss of Cold Feed
In certain columns, such as stabilizers and product strippers, a
significant amount of the reboiler duty is used to heat up the bottom
product stream. Loss of feed to such columns will result in the bulk of
the reboiler duty being available to generate relief vapor. Therefore, for
those type columns, particular attention must be paid to cases where
feed may be disrupted.
o
Reflux Failure
Loss of reflux frequently results in a significant relief load due to
subsequent flooding of the overhead condenser. Flooding is assumed
to occur if there is inadequate surge volume in the overhead
accumulator to provide enough time (10 minutes) for operator
response before the accumulator fills with liquid. When the overhead
reflux system has one turbine and one motor driven pump, assume
that the operating pump is the one that results in the maximum relief
load for each contingency (this may result in different drivers assumed
for different contingencies). Credit for autostart of the spare
pump/driver shall be considered only as part of the total flare load
analysis and will not be credited towards the relieving capacity required
for an individual tower.
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o
Pumparound Failure
Loss of circulation through a pumparound circuit results in the complete
loss of cooling from that loop, which is accounted for in the
corresponding upset heat balance. Consideration should be given for
the possibility that the loss of that loop may result in drying up of the
trays above that point and possible subsequent loss of other
pumparound circuits. Such assumptions should be reviewed with plant
operations.
o
Loss of Quench
In some columns, cooling is provided by introducing an external cool
liquid into the column for direct contact cooling. If this quench liquid
flow is lost, a relief requirement can result as the hot vapors move up
the column and vaporize top tray liquid. In some cases the load can be
large enough that the peak vapor rate calculated using the
assumptions given in paragraph 4.3.3.1 can only be sustained for a
short period before upper tray liquid is dissipated. After tray liquid is
vaporized a longer term sustainable load consisting of the hot incoming
vapors may continue to be relieved.
o
Utility Failure - (Power, Steam, C.W., Inst. Air)
Loss of utility systems are analyzed based on the resultant
consequences for the system, possibly resulting in multiple impacts
such as reflux failure, loss of condenser capacity, blocked outlets, etc.
Typically loss of power and/or loss of cooling water result in the largest
relief loads. Loss of steam is usually only critical where major turbine
drivers are impacted. Loss of instrument air results in control valves
assuming their fail-safe positions, which normally will not result in relief.
In evaluating power failure modes, it must be considered that the
condition resulting in the largest tower relief load may be different than
the failure mode that results in the largest combined flare load.
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o
Control Failure
Any one control device should be assumed to move to either the full
open or full closed position, regardless of the normal failure position.
This includes air fan louvers, which when closed will be assumed to
result in total loss of cooling duty (for the portion of the air cooler
covered by the set of louvers controlled by one signal).
o
Abnormal Heat Input
In the event of control failure on the heat input to the reboiler, it is
possible to generate vapors in excess of the condensing capacity of
the system. The unbalanced heat input in this case is calculated based
on the maximum reboiler heat input. Credit normally can be taken in
this case for the additional condenser capacity at elevated pressure,
however, since this case is usually not controlling, normal overhead
condenser duty is usually used to calculate a conservative load for this
case.
o
Absorbent Flow Failure
For lean oil absorption, no relief requirement generally results from
lean oil failure. However, if a large quantity of inlet vapor is removed in
a unit, loss of absorbent could cause a pressure rise to relief pressure,
since the downstream system may not be adequate to handle the
increased flow. The subsequent impact on downstream systems of the
additional flow and composition must be considered in the analysis.
o
Blocked Outlet
Blockage of feed, reflux, pumparound and quench streams will result in
the consequences already discussed for those failure conditions.
Blockage of the bottom stream usually does not have a relief impact,
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since there is normally a large surge volume in the bottom of columns
so that operations has enough time to respond to the event well before
relief occurs. Blockage of the net liquid overhead stream only requires
relief if the operator does not have time to respond to the drum high
level alarm before flooding the condenser. Isolation of the net vapor
overhead line will cause relief of those vapors. If those vapors are
relieved from the column instead of the accumulator, and/or assuming
the condition persists, accumulation of non-condensables is possible.
This impact is described in the next section.
o
Accumulation of Non-Condensables
Non-condensables do not accumulate under normal conditions since
they are released with the process vapor streams. However, with
certain piping configurations, it is possible for non-condensables to
accumulate to the point that a condenser is blocked. Closure of an
automatic vent valve can have the same impact.
o
Loss of Heat in Series Fractionation System
In series fractionation, loss of heat input to one column can cause light
ends to enter the bottoms stream from that column and overpressure
the following column. Under this condition, the overhead load of the
second column may consist of its normal vapor load, plus the light ends
from the first column. If this added load cannot be condensed, then the
relief will occur from the downstream column. It is also possible that
accumulation of non-condensables could result, causing loss of cooling
capacity. Relief of the additional light ends could then be additive with
the load due to loss of cooling.
o
Fire
During a fire, all feed and output streams to and from the fire affected
equipment and all internal heat sources within the process are
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assumed to have ceased. Calculation of the system fire case load will
follow normal fire case criteria.
o
Exchanger Tube Rupture
Assuming the conditions for tube rupture apply and the process side
has the lower design pressure, flow across the tube break represents
an input of both heat and material to the column. If steam or a heat
transfer fluid is used as the heating medium, the availability of the high
pressure stream can sometimes be limited by the normal position of an
inlet control valve or system hydraulics. If a process stream with
limited flow availability is used as the heating medium, the calculated
relief load is often not significant. Flow through the break is calculated
by normal procedures for a tube rupture and then checked to see if the
flow can be sustained or if it is restricted by a control valve. Assume
any fluid which enters the column reaches thermal equilibrium with the
liquid in the column at that location and that any remaining hot side
fluid exits the exchanger at the normal exit temperature. Recalculate
the thermal duty of the exchanger on that basis. Complete the upset
heat balance to determine the relief load (if any). If the heat medium is
steam or a fluid which flashes at column conditions, the vapor portion
of the incoming fluid will have to be relieved in addition to the load
calculated by the heat balance, if the vapor/steam cannot be
condensed.
4.3.5.5
Special Considerations
While the load calculation method presented provides a reasonable,
conservative and consistent approach for estimating relief rates for most
columns, it is recognized that static calculation methods have inherent
limitations. The method provided is based on the assumption that the
column internal composition profile remains the same throughout the upset
event and that the liquid on each tray is at its bubble point. These
assumptions are most closely met for a column with a large number of trays
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separating similar components (similar boiling points). Engineering judgment
should always be used in evaluating relief loads for columns which operate
much differently than supposed in these assumptions.
Following are some examples where special evaluation techniques are
required.
o
Steam Stripped Fractionators
Steam stripping is typically used in crude towers, for example. When
calculating the bubble point temperature and enthalpy of the individual
product streams, use the COX chart to extrapolate to relief conditions.
Use of simulation data, unless compensated for steam partial pressure,
can lead to erroneously high stream temperatures and lower calculated
relief loads.
o
Bottoms Stream Heat Recovery in Non-Reboiled Fractionators
In columns such as Crude Towers, where preheated feed is flashed
into the bottoms of the column, subsequent heat recovery from the
bottoms stream should be estimated using the normal column bottoms
temperature. At the elevated relieving pressure, more light ends go
with the bottoms stream and the assumption of normal temperature for
heat integration purposes is usually conservative. (Note: this does not
influence the use of the COX chart in determining a different bottoms
temperature for overall heat balance purposes as described above).
o
Reduced Heat Input in Non-Reboiled Fractionators
In columns such as Crude Towers, where preheated feed is flashed
into the bottom at the column, it is possible to underestimate the relief
load when feed preheat is substantially reduced from normal
conditions. Results from one dynamic simulation for a Crude Tower,
for example, indicated that the change (reduction) in relief load from full
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heater duty to half heater duty was only half of that predicted by the
standard calculation method. This result is attributed to the sudden
change in temperature of the feed, while significantly higher liquid
temperatures still are present in the bottom of the column. As a result,
for a period of time the vapors are reheated in the bottom of the
column and subsequently vaporize an increased amount of liquid in the
upper section of the column. This transient effect is not possible to
predict using static calculation procedures. Unless dynamic analysis is
applied for the specific column being analyzed, it is recommended that
relief loads for feed heat reduction in this type of column be calculated
using normal procedures, except that the results of this one dynamic
study will apply as follows:
Relief load @ reduced feed preheat duty = ½ [Calculated relief load @
full feed preheat duty + calculated relief load @ reduced feed preheat
duty].
This is intended to provide a conservative relief load estimate.
o
Strippers for Absorption Systems
Stripping columns which remove absorbed components from a solvent
stream require special consideration, since typically the solvent is not
vaporized in the column. Using the normal calculation procedures can
erroneously indicate that no vapor is generated due to the large
amount of energy required to heat the bottom solvent stream. A
reasonable estimate of the quantity of vapor flowing to the overheat
condenser can frequently be obtained by increasing the feed enthalpy
by the normal reboiler heat input and flashing the feed stream at relief
pressure.
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4.4
REACTOR LOOPS
4.4.1
System Description
Reactor recycle loops discussed in this section are defined as recycled, fixed-bed,
pressurized reactor loops typical of those used as hydrotreaters, hydrocrackers, and
catalytic reformers. These processes operate at moderate to high temperatures
and pressures in a recirculating hydrogen atmosphere. Depending on the reactions
being carried out, the process will be a net generator or consumer of hydrogen and
correspondingly will either absorb or evolve heat. Many reactors operate with totally
vaporized feed, although reactors processing heavier feedstocks generally operate
in two-phase flow. The reactor loop is usually protected by one or more pressure
relief valves located at the low pressure side of the loop, usually at or near the
product separator.
4.4.1.1
Process Flow
A typical, simple loop consists of the reactor(s), charge heater,
feed/effluent exchanger, product coolers, product separator and the
recycle compressor, with the associated piping and controls. Key
operating variables include temperature, pressure, space velocity, and gasto-oil ratio. Permissible ranges for these variables are defined during
process design.
4.4.1.2
Start-of-Run and End-of-Run Conditions
Catalytic reactors have a finite run length, at the end of which the catalyst
must be regenerated or replaced. As the cycle proceeds, increases in
reactor temperature are used to maintain conversion as catalyst
deactivation proceeds. Product yields decline and light ends make
increases, due to reduced selectivity. A run is ended when operating
temperature limits are reached or yield loss is excessive.
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End-of-run (EOR) conditions are generally more severe than start-of-run
(SOR). EOR conditions typically result in higher operating temperatures,
higher molecular weight and higher loop pressure drop. Each of these
tends to increase pressure relief valve sizing or set point so that EOR
conditions are typically assumed to be controlling for relief system design
purposes.
4.4.1.3
Reaction Process Characteristics
Exothermic processes (such as hydrotreaters and hydrocrackers) are
characterized by:
o
o
o
Net consumption of hydrogen.
Heat release, which tends to raise the system temperature.
A decreased number of moles, which if reaction continues in isolation
will outweigh increased temperature and cause system pressure to
decline.
Endothermic processes (such as catalytic reforming) are characterized by:
o
o
o
Net generation of hydrogen.
Absorption of heat, which tends to lower the system temperature.
An increase in the number of moles in the system, which if continued
in isolation will increase overall system pressure.
4.4.1.4
Alternate Operation Modes
During the course of startup, shutdown, and catalyst regeneration, the
reactor loop may be filled with different fluids at conditions of temperature
and pressure different from normal operation. Generally, operating
conditions for these circumstances are selected to be within the capability
of the equipment designed for normal operation. Therefore, alternate
operations can usually be neglected in terms of reactor loop pressure relief
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system design. However, special or dedicated systems only used during
these occasions (such as regeneration skids) must still be evaluated.
4.4.2
Causes of Overpressure
The potential relief causes affecting reactor loops may be grouped into three main
categories, as follows:
o
Operating Failures
-
Loss of Feed
Loss of Effluent Cooling
Loss of Quench
Recycle Compressor Failure
Utility Failure: power, steam, cooling water, instrument air
Control Failure
Blocked Exits
Abnormal Heat Input
o
Compositional Changes
Change in Feed Composition
Chemical Reaction
o
Other Conditions
-
4.4.3
Fire
Heat and Material Balance Considerations
4.4.3.1
Basic Assumptions for Operational Upsets
o
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A basic assumption for operational upsets is that the normal reactor
effluent flow continues and that the recycle compressor continues to
circulate vapor.
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4.4.3.2
o
Even if liquid feed is lost, normal effluent flow is assumed to continue
for a short period (conservative assumption based on bulk of volume
being recycle gas).
o
Recalculate effluent exchanger duties for upset conditions based on
any changes in flow and temperature (LMTD) resulting from the upset.
o
Assuming excess vapor is generated due to loss of effluent cooling
capacity, credit may be taken for added capacity of the recycle
compressor at relief conditions and increased flow through the offgas
valve in its normal position. Any surplus vapor (above normal flow)
after deducting these credits, must be relieved.
o
Where licenser input is available, it can be used to supplement or
modify the approach based on experience.
Reactor Yields
Reactor conversion is not changed by upset conditions for the short term
relieving event. Normal effluent flow and composition continue.
Due to large heat content in reactor, outlet temperature changes slowly. Use
normal EOR reactor exit temperature for most upsets. Cases involving
excess heat of reaction, loss of quench, or abnormal heat input usually do
not impact relief system sizing.
4.4.3.3
Condensation Curves
Increased pressure during relief does not significantly impact the heating and
cooling curves. Therefore, it is acceptable to check performance at either
normal conditions based on existing cooling curves, or at the higher
pressure, if a more precise calculation is required or preferred.
When an upset results in reduced condensing capacity, the molecular weight
of the vapor flowing to the separator is based on the revised performance of
the effluent exchange train and the condensation curve. This is the
composition of the material to be relieved.
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4.4.4
Pressure Profiles
It is not economical to design all the equipment in a reactor loop for a single high
pressure, therefore the dynamic pressure drop around the loop is taken into account
in establishing equipment design pressures. In order to confirm the adequacy of the
pressure relief protection, the analysis of each operational failure must likewise
consider the associated pressure profile.
4.4.4.1
Operating Pressure Profiles
The pressure in a reactor loop is automatically controlled at a single point,
typically at the product separator, and is usually held constant during the
run. The operating pressure profile for the loop is established simply by
successively adding the pressure drop associated with each item or line
segment, starting with the pressure at the control point.
Pressure drop tends to increase as the operating cycle proceeds. When
dealing with a new unit, the pressure profiles for both start-of-run (SOR)
and end-of-run (EOR) operations should be available as part of the
process design data. For an existing unit, it is preferable to use observed
operating data to the greatest possible extent.
4.4.4.2
Design Pressure Profile
As a minimum, the design pressure of the vessel on which the pressure
relief valve is located must be high enough to give an adequate margin
over normal operating pressure. Once the separator design pressure is
established, it becomes the datum for setting all the other design
pressures in the loop, using increments which correspond to the governing
EOR operating pressure profile.
In a new unit, there may be considerable incentive to provide additional
margin in the separator pressure design in order to allow contingency for
future capacity, or to minimize the impact of certain upset cases. For
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example, in API RP-520, Part I, it is suggested that the separator should
be designed for 105% of the loop settle-out pressure. The latter is a
conservative approach which may or may not be justified. However, it
must be recognized that once the design pressure profile is established, it
will impose operating limitations which cannot be exceeded without putting
the adequacy of the pressure relief system in question.
4.4.4.3
Relieving Pressure Profile
The relieving pressure profile represents pressure levels around the loop
under conditions when the pressure relief valve is expected to operate.
The relieving pressure profile must be examined versus equipment MAWP
both when the inlet to the pressure relief valve is just at set pressure and at
fully accumulated relief conditions.
To get an approximation of these profiles one can assume the EOR
pressure profile is simply shifted upward by an increment equal to set
pressure minus the normal operating pressure at the pressure relief valve
location (incipient relief), or accumulated relief pressure minus normal
operating pressure (fully accumulated relief). This approach assumes that
pressure drops through individual equipment items are unchanged at relief
conditions. This is a directionally conservative assumption in a number of
relief cases as, at constant mass flows, the higher gas densities at
relieving pressure will tend to reduce the pressure drop. EOR pressure
drops should be adjusted as necessary in those segments which are
affected, if there is a significant change in mass flow, a significant vapor
density change due to the pressure rise, or a significant temperature/phase
change in a portion of the loop due to loss of cooling.
Different relief contingencies may affect pressure drops in one portion of
the loop and not another, so that it may be necessary to develop profiles at
accumulated relief pressure for more than one case. Usually, no more
than one or two major cases need to be considered.
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In a case where the design pressure for any equipment item would be
exceeded before the pressure relief valve opened, the proposed protection
would not be adequate and one or more of the following remedial
measures must be implemented to achieve a viable design:
o
Lower the set pressure of the pressure relief valve, if possible. If the
margin between operating and set pressure is already at a minimum
for a conventional spring-opposed pressure relief valve, consider a
pilot-operated pressure relief valve as a substitute.
4.4.4.4
o
Under administrative controls, impose limits on the allowable unit
throughput to reduce the slope of the pressure profile.
o
Rerate the non-compliant items for higher design pressure.
Settle-Out Pressure
In compressible flow applications with significant dynamic pressure losses,
“settle-out-pressure” refers to the pressure level at which the system will
equilibrate if gas flow is suddenly lost. In a reactor loop, settle-out will
result from any situation in which the recycle compressor is tripped, as
liquid-phase pressure drops alone are usually negligible. Depending on
the steepness of the operating pressure profile and the available margin
between normal separator operating pressure and design pressure, settleout may even result in an over pressure situation in which the pressure
relief valve opens briefly. Settle-out pressure can be estimated by dividing
the loop into sections, calculating the mass of vapor in each segment,
estimating a weighted average vapor temperature and molecular weight,
then calculating the pressure based on total system volume.
4.4.5
Pressure Relief and Depressuring Facilities
There are a number of key considerations related to the location and setting of the
primary loop pressure relief valves, which must be carefully observed to insure that
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code requirements are satisfied. Where isolation valves are present in the circuit,
supplemental pressure relief valves may be required to protect individual items of
equipment. These and other aspects of loop protection are discussed in this
section, including the use of pilot operated pressure relief valves and the role of
depressuring facilities.
4.4.5.1
Code Criteria
The three code-based requirements for protecting a system of
interconnected vessels with a single pressure relief valve installation that
are of critical importance to reactor loop relief protection are as follows:
1.
The set pressure of the pressure relief valve must be at or below the
lowest MAWP of any equipment item in the reactor loop which it is
intended to protect. If multiple pressure relief valves are provided,
the set pressure of at least one valve must meet this requirement.
Other valves may not be set higher than 105% of the lowest MAWP
(with the exception of a pilot-operated pressure relief valve with the
sensing point located on upstream equipment).
2.
The operating pressure in any equipment item in the loop must never
exceed its MAWP when the pressure relief valve protecting the
system is not discharging.
3.
When the pressure relief valve protecting the system is discharging,
the accumulated pressure in any equipment item must not exceed
allowable code limits for either operating or fire contingencies, i.e.
110%, 116%, or 121% of that item MAWP, as applicable.
4.4.5.2
Location of Pressure Relief Valves
The primary pressure relief valve should be located on or near the product
separator, since it should be set to protect the lowest design pressure
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equipment in the loop. There are other practical benefits with this
arrangement:
o
At this point in the system, the effluent has been fully cooled. The
pressure relief valve does not see the temperature extremes of the
reactor, and the flowing volumes have been significantly reduced by
cooling and condensation. Relief vapors at this point are often safe
for atmospheric discharge (see Section 9.0).
o
4.4.5.3
The pressure relief valve is downstream of the main liquid/vapor
separator, so that the chances of liquid or mixed phase relief are
minimized (but usually not eliminated).
Presence of Block Valves in the Loop
To satisfy code requirements, there must not be any means of blocking
any equipment item from access to the loop’s pressure relief valves. This
is generally not a problem, as isolation valves are not provided in the loop.
If block valves are present, the following guidelines should be observed:
o
If the block valves are installed strictly for maintenance, e.g., for
shutdown isolation of an air cooler, they should be locked open
during operation.
o
If the valves are installed to take an item out of service and bypass it
during normal operation, administrative controls must be imposed to
ensure that item is vented and drained whenever it is isolated. No
additional pressure relief valve is required. Key interlock systems are
sometimes required by local authorities for this type of installation.
o
If the item in question cannot be vented when it is isolated, e.g., a
standby reactor left under a gas blanket, then an individual fire
pressure relief valve must be provided for that item.
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Even though overall protection is provided by the separator pressure relief
valve, there may be circumstances other than those just described in which
individual pressure relief valves are needed for some items within the loop.
For example, a pressure relief valve may be needed at the recycle
compressor discharge.
4.4.5.4
Pilot Operated Pressure Relief Valve Applications
In some reactor loop evaluations, in particular when considering the relief
protection of an existing unit running at a higher than design throughput,
the use of pilot operated pressure relief valves may be considered in order
to permit operating pressure to approach design pressure more closely.
Either a direct-sensing or a remote-sensing pilot might be needed,
depending on the application. In any case a pilot operated pressure valve
is proposed, both the selection of the pilot operated pressure relief valve
and the details of its installation must be reviewed with the valve
manufacturer.
4.4.5.5
Depressuring
Reactor loops, particularly those which operate at higher pressure, are
often provided with automatic or operator-initiated emergency depressuring
facilities. These are provided for a number of reasons, including:
o
Reduction in contained pressure to minimize the possibility of
equipment failure through overheating during fire exposure.
o
Reduction in hydrogen partial pressure to “kill” exothermic reactions
in response to a major loss of circulation, such as recycle compressor
failure.
o
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Generation of induced flow to protect heaters and sweep catalyst
beds, also in response to a loss of recycle circulation.
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o
Safe venting of contents to remove the source of high pressure fluid
in the event of a major leak.
The ability to rapidly depressure a reactor loop is thus an essential safety
feature, and provides a function which a pressure relief valve cannot.
Whether automatic or manually activated, however, there is no situation in
which a relieving case is either eliminated or reduced in magnitude by the
presence of depressuring facilities.
4.4.6
Maximum Capacity
If it is determined that the relief system will be designed for the maximum
capacity, careful assessment of equipment limitations and operating
constraints is required. If a new unit is being considered, assumptions may
have to be relied upon to a greater degree to identify probable bottlenecks.
The maximum capacity of a unit may be determined by any one of a number
of factors, including:
o
Pump or compressor limitations: Inability to increase feed rate, or to
maintain satisfactory recycle ratios beyond a certain point.
o
Heat transfer limitations: Shortage of surface in feed/effluent
exchanger or product cooling.
o
Reactor limitations: Cannot operate satisfactorily below a given
catalyst-to-oil ratio. (Note: this may not be an “absolute” bottleneck,
as an improved catalyst or an operating decision to accept reduced
catalyst life in exchange for increased throughput may permit higher
capabilities to be attained.)
Regardless of equipment capabilities, each reactor loop also has a hydraulic
throughput limit which cannot be exceeded. The hydraulic limit is reached
when the relieving pressure profile for a given throughput coincides with the
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design pressure profile at one or more points in the loop, and it is not
possible to create any additional margin by lowering the pressure relief valve
set pressure.
4.4.7
Determination of Relief Loads
Relief loads for reactor loops are calculated using the procedures in this
section. A sample calculation is provided in Appendix B-2.
One of the basic assumptions in Section 4.4.1 is that the recycle compressor
operation continues through the operational upset. With this consideration,
performance of the effluent exchange train is recalculated to obtain the total
vapor entering the product separator under the upset conditions. With loss of
cooling capacity, the vapor will contain product that would normally condense.
Vapor flow rate, temperature and molecular weight will all increase relative to
normal flow, causing an overpressure event to the extent that the excess vapor
exceeds normal outflow capacity. Separator vapor which continues to flow to
the compressor suction is deducted from the total incoming vapor. Remaining
vapor flow must leave the loop, either via the normal vent valve or the relief
valve.
Recalculate the recycle compressor suction flow based on the normal
operating speed (constant volume). Check horsepower limitations using
design efficiencies and the estimated differential corresponding to the relieving
pressure profile. If power is limiting, take only a pro-rated credit based on the
available horsepower. As a first approximation, the relieving pressure profile
can be assumed to be the same as the operating pressure profile, adjusted
upward uniformly by the incremental pressure difference between accumulated
relief pressure minus normal operating pressure. Adjustments can be made
where there are significant changes in mass or volume flow, such as for loss of
feed (in the affected segments) or for significant temperature changes.
These bases are intended to be conservative, so that there is no need to refer
to compressor curves or to converge on a new steady state solution based on
recycle gas gradually becoming heavier over time.
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If the net off gas from the loop is taken from the discharge of the compressor,
only the flow through the compressor can be credited as a deduction against
the total separator outlet flow.
Following are the impacts of the major contingencies to relief loads:
4.4.7.1
Loss of Feed
Loss of liquid feed to a reactor loop is analyzed based on the transient loss
of cooling to the feed/effluent exchange train. Effluent temperature out of
the feed/ effluent exchange is based on recalculated exchanger performance
without liquid feed. Downstream effluent exchangers are similarly rerated for
the higher flow and temperature exiting the feed/effluent-exchange train. Net
relief load is calculated as indicated in Section 4.4.3.
4.4.7.2
Loss of Effluent Cooling
This is the generalized case which can, for example, include high
temperature feed, loss of cooling water or loss of air cooler fans. A similar
approach is used to that described for loss of feed for the affected exchanger
and the downstream exchange train. In cases where loss of cooling
prevents condensation and there is no other mechanism for removal of
products (such as net gas compression), then the relief rate must equal the
feed and makeup rates on a mass basis.
4.4.7.3
Loss of Quench
Normally quench control valves are designed to fail open and loss of only
one quench stream is considered at one time. This impact can be evaluated
conservatively as a loss of the final quench point, by using the corresponding
reactor effluent temperature and the normal outlet flow to calculate a relief
load similar to the abnormal heat input case.
4.4.7.4
Recycle Compressor Failure
Loss of the recycle compressor effectively stops reactor loop circulation in
most cases since recycle flow represents a significant percentage of total
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circulation. In high pressure systems, automatic or manual depressuring of
the loop is often initiated on loss of the recycle compressor to avoid potential
negative consequences from loss of circulation. However, the presence of
depressuring capability does not eliminate the need for pressure relief
protection.
Excessive vapor temperature will cause major damage to a compressor.
Loss of cooling capacity will cause the vapor temperature to increase. High
temperature shutdown trips are normally used to prevent compressor
damage. In the absence of more specific information, it can generally be
assumed that if loss of cooling causes the compressor inlet temperature to
rise above 392 °F (200 °C), then the compressor will be shut down.
If loop settle-out pressure exceeds the separator relief valve set pressure
plus allowable accumulation, there is a viable relief case on compressor
failure. Flow reaching the separator at this pressure will have to be relieved.
Precise calculation of the relief load in this case is difficult without using
dynamic simulation, however a reasonably conservative estimate can be
made as follows:
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o
Using the normal operating pressure profile, locate the point where
the pressure is approximately equal to the settle-out pressure.
Determine the total flow rate, pressure drop and average mixed
phase density between that point and the product separator (under
normal conditions).
o
Assuming this pressure remains about the same after the
compressor trips, estimate the flow to the separator at accumulated
relief conditions based on the new pressure drop available. The
peak relief load will be the vapor portion of the estimated separator
inflow less any credit for vent gas withdrawal. If make-up gas
continues, the relief load will eventually approach the make-up gas
flow less credits. Make-up gas flow should not be added to the
transient relief load, which is dominated by system hydraulics.
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4.4.7.5
Utility Failure (Power, Steam, Cooling Water, Instrument Air)
The impacts of utility failure result from subsequent equipment outage, flow
loss, cooling loss and other consequences of local or plant-wide outages.
Power or steam failure can result in loss of feed, loss of the recycle
compressor, or loss of air cooling, depending on whether equipment is motor
or turbine-driven. Instrument air failure can result in loss of feed and blocked
vapor/liquid outflows. Cooling water failure can cause loss of effluent cooling
and loss of the recycle compressor (for condensing turbine drive). Resultant
relief loads are analyzed as discussed in other sections of this write-up for
the particular loss which occurs.
4.4.7.6
Control Failure
Individual failure of controls must be reviewed along with loss of instrument
air. Total failure due to loss of instrument air will cause control valves to
move to their fail-safe positions. This typically includes isolation of the feed
and vapor/liquid product streams, cut-off of fuel to the fired heater and
continued operation of the recycle compressor.
4.4.7.7
Blocked Outlet
Blockage of the vapor exit line will result in overpressure since vapors will
accumulate in the system. In hydrogen-producing systems, the relief load is
equivalent to the maximum net make gas. In hydrogen consuming systems
it is the maximum net make-up gas rate less the reactor purge or bleed
stream (no deduction for chemical consumption or solution loss, since they
may not be occurring).
A blocked product liquid line will cause buildup of liquid in the separator and
eventually can result in carryover of liquid to the compressor suction drum,
resulting in a high level shutdown. Added pressure buildup can result in
relief of some liquid through the separator pressure relief valve, if fill time is
less than that required for operator response. While the impact on relief
valve sizing must be considered, generally it is concluded that relief valve
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sizing is not impacted, but that provisions for liquid release into the relief
header must be made.
Loss of separator bottoms liquid flow can also cause relief if the liquid is
normally exchanged with reactor effluent. In that case, the normal approach
in evaluating loss of cooling is used.
4.4.7.8
Abnormal Heat Input
This case usually arises from a fired heater going to maximum firing
(assumed to be 125% of design heat input), and is treated like a loss of
cooling case, except that the reactor effluent temperature is increased by the
additional duty. Effect on reactor performance is neglected. Typically, the
calculated relief load for this contingency is small due to the small amount of
preheat duty added for exothermic reactions and the use of multiple reactors
and heaters for endothermic systems (only one heater assumed to be
affected).
Loss of feed could be considered in this category if the analysis is extended
past the feed/effluent exchange train into the heater(s) and reactor(s).
However, this contingency is normally not considered because the recycle
gas alone can only absorb a fraction of the normal heater duty and this
thermal effect would not be additive to the lost cooling considered in the loss
of feed case (since reactors will already be exhausted of reactant). Further,
continuing to fire the heater at normal loads would result in tube overheating
and failure.
4.4.7.9
Change in Feed Composition
Certain feedstocks are unusually reactive (for example hydrotreating of
cracked stock) or result in a significant energy release and usually represent
only a fraction of the total unit feed. Undiluted charge of the feed can result
in localized overheating if fed at its normal rate. This rarely results in a relief
case. However, if the reactive feed can be unexpectedly charged at full unit
feed rate, when normally the feed is diluted by less reactive material, the
consequences can be catastrophic. Generally, overpressure protection is
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not sufficient to protect the system in these circumstances, and other
safeguards must be put in place to prevent the situation from occurring.
4.4.7.10 Chemical Reaction
In general the reactions considered in loops will tend to generate minor relief
loads, if any. Concerns in exothermic systems are primarily related to
overheating which is not helped by overpressure protection. Endothermic
hydrogen-producing reactors are limited by the ability of the system to supply
reactor heat. While the physical system can supply heat for a short period of
time, it is unlikely that any case generated would require relief equal to the
normal net make gas which is the case already considered under blocked
exit.
4.4.7.11
Fire
During a fire, the unit is assumed to be shut down with normal contents
intact. For vessels and exchangers normally in all vapor service, ignore the
fire heat input, since vapor expansion will have a minimal contribution to the
total load. Concern for the integrity of these items of equipment may result in
depressurizing capability being installed. It is acceptable to simplify
characterization of the liquid contents of each item by assuming that the
entire heat input is used to vaporize normal separator liquid.
4.5
LIQUID FILLED SYSTEMS
4.5.1
Blocked Discharge
Liquid filled systems present application problems which are different from those
found in vapor systems. Among the characteristics that make handling of design of
pressure relief systems for liquid-filled systems different from vapor systems are:
o
Liquid filled systems have very little capacitance, so system response to over
pressure causing upsets and to operation of pressure relief devices is very
rapid and much more prone to instability.
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o
Liquid relief valves have operating characteristics which are different than
pressure relief valves in vapor services.
o
Liquid relief services often occur in systems with centrifugal pumps. The
required relief load is a function of the pump characteristics, system
hydraulics, fluid properties and operating conditions on the suction side of
the pump.
o
4.5.2
Liquid relief services present greater difficulties in disposing of relief streams.
Thermal Relief
Consideration shall be given to liquid filled piping or equipment which may be
closed-in and exposed to a heat source other than fire. The resulting hydraulic
expansion requires relief of a small amount of liquid. If the temperature of the heat
source is high enough, vaporization of the liquid at relieving pressure may occur.
Blocking in may be caused by any combination of closing of block valves, check
valves and control valves. Heat sources to be considered include the following:
o
Steam tracing
o
Ambient temperature heat input to equipment filled with refrigerated liquid.
o
Solar radiant heat input
o
Hot side fluid of heat exchanger
Normally only very low relief capacities are needed and therefore relief valves are
normally selected as a nominal minimum size of 3/4” X 1” NPS (Nominal Pipe Size)
without relief rate or sizing calculations. However, where it is possible to have a
very high heat input rate to the blocked-in liquid filled system, calculate the liquid
relief rate basis using the formula in Appendix C of API RP 520 Part I, 5th Edition.
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Two general applications for which thermal relieving devices larger than a 3/4-inch X
1-inch nominal pipe size (NPS 3/4” X NPS 1) valve might be required are long
pipelines of large diameter in uninsulated aboveground installations and large
vessels or exchangers operating liquid full. Long pipelines may be blocked in at or
below ambient temperature, and the effect of solar radiation will raise the
temperature at a calculable rate. If the contained volume and thermal-expansion
coefficient for the fluid are known, a relieving capacity requirement can be
calculated.
4.6.
MECHANICAL EQUIPMENT
4.6.1
Pumps
Pumped systems need to be protected from overpressure if the maximum discharge
pressure, computed from the following two cases exceeds the pump casing design
pressure, pump discharge piping design pressure, or the MAWP of any vessels in
the discharge piping system.
o
Suction pressure at normal operating pressure, discharge pressure at shutoff
head at trip speed of the driver. Depending on the basis for system design,
either the installed or maximum sized impeller is considered.
o
Suction pressure at the relief pressure of upstream equipment, including
accumulation, and the pump at process design normal flow. Only upstream
relief conditions that are credible during system operation are considered.
This excludes fire relief conditions.
If a centrifugal pump has been constructed in accordance with API Standard
610, the maximum allowable casing working pressure must be at least equal
to the maximum discharge pressure, and pressure relief protection for the
pump itself is not required. If possible, however, the designer should
independently confirm the adequacy of the casing design pressure,
particularly in the case of existing pumps where service conditions may have
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changed significantly. Even where the pump does not require over pressure
protection, the vessels downstream of the pump may.
4.6.2
Compressors
Centrifugal compressors must be protected with pressure relief devices in all cases
where the maximum discharge pressure, determined as in the following paragraph,
exceeds the maximum casing working pressure defined by the compressor vendor.
Maximum discharge pressure of a compressor correspond to the intersection of the
100% speed line with the surge line. In the case of turbine-driven compressors, the
105% speed line is used. If blocking in the compressor will cause suction pressure
to rise, this must be taken into account. The effects of higher than normal molecular
weight or variations in suction temperature also need to be considered.
Although a discharge pressure relief valve will help to maintain flow through the
machine in the event of a blocked discharge, it is not a substitute for a properly
designed anti-surge system.
The ANSI B19.3 Safety Standard for Compressors indicates that gas compressors,
drivers and their auxiliaries shall be protected by a pressure-relieving device or
devices to prevent the pressure in any element of the system from exceeding 110
percent of its maximum allowable working pressure (MAWP). Exceptions to this
requirement may be made in systems in which the only possibility for pressure
exceeding 110 percent of MAWP relates to accidental closure of block valves,
provided such valves are intended only for isolating equipment that is shut down.
The ANSI standard also states that to minimize leakage from pressure relief valves,
the set pressure of valves should preferably be a minimum of 10 percent or 15 psi
(1.03 bar, 1.05 kg/cm2) (whichever is greater) above the intended operating
pressure at the valve inlet.
The relief valves shall be sized in accordance with ASME Code. Where the set
pressure is below the MAWP of the protected equipment, the MAWP may be
substituted for set pressure in the code calculation.
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Reciprocating compressors require special considerations because of the pressure
pulsation. Relief valve manufacturers generally state that the differential between
operating and set pressures may need to be greater than the 10 percent normally
recommended. It is always good practice to set the valve as high above the normal
discharge pressure as possible.
4.6.3
Mechanical Driver Considerations
o
Steam Turbines
Noncondensing Turbine - A steam turbine designed to operate with an
exhaust steam pressure equal to, or greater than, atmospheric pressure.
NEMA Standards Publication No. SM23 should be consulted for detailed
requirements of both the sentinel warning valve and the full flow relief valve.
Full-flow relief valves shall be provided for all multi-stage steam turbines and
all single-stage steam turbines, unless the turbine outlet casing can
withstand maximum inlet steam conditions. The full-flow relief valve is part
of the piping installation which is external to the turbine and is installed
between the turbine and the first shutoff valve. The relief valve should not be
confused with the excessive-exhaust-pressure warning valve (sentinel type)
which is mounted on the turbine casing. Sentinel type safety valves for
turbine cases will be specified as part of the turbine specification.
Where a full-flow relief valve is required it shall be sized such that it will
exhaust to the atmosphere the maximum quantity of steam (as determined
by the turbine manufacturer) which will pass through the turbine nozzles
under rated initial conditions. Therefore, it is important to request this data
early to insure timely receipt of certified data from the vendor.
o
Condensing Turbine
A steam turbine designed to operate with an exhaust pressure below
atmospheric pressure. The exhaust system must be protected against
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excessive pressure by the installation of an atmospheric type safety valve,
usually on the surface condenser, capable of passing all of the steam which
may reach the turbine exhaust without the pressure rising to a value greater
2
than 10 psig (0.7 barg, 0.7 kg/cm ) + 10 percent accumulation. The size of
the atmospheric safety valves recommended for use with condensing
turbines is shown in Table 4.3. Normally the “For Protection” application is to
be used.
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TABLE 4.3
CONDENSING TURBINES - ATMOSPHERIC SAFETY VALVE SIZES
Note:
Maximum Steam flow, kg/hr
For Protection
(NPS)
Up to 3,402
3,403 to 5,352
5,353 to 7,711
7,112 to 9,072
9,073 to 10,478
10,479 to 13,698
13,699 to 17,327
17,328 to 20,412
20,413 to 21,409
21,410 to 28,123
28,124 to 30,844
30,845 to 37,194
37,195 to 48,081
48,082 to 54,431
54,432 to 77,110
77,111 to 113,398
113,399 to 172,365
172,366 to 249,475
6
8
8
8
10
10
12
12
14
14
16
16
18
18
20
24
30
36
For Maximum NonCondensing
Operation
(NPS)
8
10
12
14
14
16
18
20
20
24
24
30
30
The sizes listed “For Protection” are normally used under ordinary condensing operation
and are for general reference only. If it is desired to operate the turbine temporarily
noncondensing at its maximum noncondensing capacity, the sizes listed under “For
Maximum Non-Condensing Operation” should be used. Actual design conditions, i.e.,
flow, relieving pressure, should be established by the user and condenser
manufacturer. The valve relieving capacity and design should be certified by the valve
supplier.
Reference:
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Heat Exchange Institute, Standards for Direct Contact Barometric and low level
Condensers - 1970
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4.7
HEAT EXCHANGER TUBE RUPTURE
ASME Section VIII Code, Division 1, paragraph UG-133(d) requires protection of
exchangers against internal failure. As indicated in API RP 521 (Third Edition), Section
3.18.1, this ASME code section defines a broad problem, but does not provide specific
direction regarding the extent of the failure, the magnitude of the relieving capacity
required, or how to properly relieve the system. API RP 521 provides the following
philosophy, which has been reviewed and approved as good engineering practice by API
member companies, as a solution:
The likely mode of exchanger internal failure is considered to be a tube leak. During normal
operation, this leakage would typically be absorbed without overpressure by the low
pressure side. Further, the low pressure side is typically protected by a relief device
somewhere along its flow path. If the low pressure side can be blocked in with the high
pressure side still pressurized, then it is necessary to provide a relief device to protect the
low pressure side. This is of great importance. There have been several reports of
incidents when this possibility was ignored:
o
High pressure liquid in an exchanger caused a rupture of the low pressure side
when the low pressure side was blocked in, resulting in loss of life.
o
A water jacket on a high pressure compressor was blocked in. Seal leakage
resulted in blowing the head off the compressor.
Experience has shown that complete failure of an exchanger tube is a remote probability. It
should be addressed more from the point of hazards analysis based on potential impact.
Guidance is provided in API RP 521 on how to address this issue with the following intent:
a)
Design Pressure Considerations
If the low pressure system hydrotest pressure is above the high pressure side
design pressure, the likelihood of losing fluid containment and the impact of loss of
containment are typically low. As a result, a relief analysis is not usually required for
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exchangers where the design pressure of the low pressure system is at least twothirds (2/3) the design pressure of the high pressure side (based on typical hydrotest
pressure at 1.5 times the design pressure). See paragraph 4.7.2
Conversely, exchangers for which the low pressure side hydrotest pressure can be
exceeded should be analyzed. If in addition, the exchanger is designed for high
2
pressure [>711 psig (49 barg, 50 kg/cm G)], the low pressure side is liquid full, and
the high pressure side contains a gas or a fluid that will flash across the rupture,
dynamic analysis of the system should be considered. See paragraph 4.7.3. If a
relief valve is installed as a critical means of protection, then the normal installation
criteria apply, including setting lift pressure at or less than design pressure and
allowing 10% accumulation maximum.
b)
Prevention of Tube Rupture
Maximum emphasis in preventing overpressure due to tube rupture should be
placed in the prevention of tube rupture itself, rather than in reliance on relief
systems to protect after the failure. This can be accomplished by giving special
attention to exchanger design/specification, inspection and maintenance. This is
particularly true for services which have special concerns or some history of
problems. The unique design features of each exchanger being evaluated need to
be considered in arriving at a design that adequately addresses overall safety.
The above philosophy is similar to considerations provided for double jeopardy or other
remotely possible modes of failure, such as loss of reactor internals resulting in hydraulic
blockage. While not normally considered as a viable relief contingency, where there is
some history of unusual failure modes they may need to be considered as part of hazard
analysis in systems where the consequences are significant should the infrequent event
occur. Addition of a relief device is just one option in mitigating such hazards.
4.7.1
Determining Required Relief Flow Rate
Assume the following basis to determine the required relieving flow rate.
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o
The tube failure is a sharp break in one tube
o
The tube failure occurs at the back side of the tubesheet
o
The high pressure fluid flows into the low pressure side of the exchanger
through both the tube stub remaining in the tube sheet and the other, longer
section of the tube. For simplification, the calculation equations provided are
conservatively based on flow through two orifices.
Calculations must account for any liquid that will flash to vapor either as a result of
the pressure reduction or because of the combined effects of pressure reduction
and vaporization if the fluid is intimately contacted by hotter material on the low
pressure side. The impact of flashing as a result of pressure reduction is calculated
using the methodology discussed later in this section. If flashing may result from
the mixing of the two fluids, the low pressure side fluid flow rate should be
recalculated at relieving pressure, and the two streams mixed adiabatically to
determine the extent of flashing induced by mixing. Then the steady state analysis
methods provided in Section 4.7.2 should be applied to arrive at the required
relieving rate.
Liquid Flow
For liquids which do not flash when they pass through the opening, the liquid flow
rate through the failure should be calculated using the following incompressible flow
equation:
W L = 2 (1891) d2C(∆PρL )0.5
4.3a (English)
2
0.5
4.3a (Metric)
2
0.5
4.3a (Metric)
W L = 2 (816) d C(∆PρL )
W L = 2 (808) d C(∆PρL )
(See Crane Technical Paper 410, Equation 3.21.
Note: Equations used here are conservatively based
on flow through two orifices.)
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Assume C = 0.6
Note: This is equivalent to a beta ratio less than about
0.2 and a Reynolds Number above 1,000.
W L = 2 (1135) d2(∆PρL )0.5
4.3b (English)
2
0.5
4.3b (Metric)
2
0.5
4.3b (Metric)
W L = 2 (490) d (∆PρL )
W L = 2 (485) d (∆PρL )
Where:
WL
=
Total liquid flow rate through both tubes, lb/hr (kg/hr)
d
=
Tube I.D., inches
P1
=
High pressure side operating pressure, psia (bara, kg/cm2 A)
P2
=
Low pressure side accumulated relief pressure (backpressure),
psia (bara, kg/cm2 A)
∆P
=
Differential pressure or pressure drop = P1 - P2 , psi (bar, kg/cm2)
ρL
=
Liquid density at flowing conditions, lb/ft3 (kg/m3)
C
=
Orifice coefficient (see Crane Technical Paper 410, page A-20)
Vapor in Critical Flow (P2 less than or equal to Pcf)
1)
Check for critical flow:
Pcf = (P1)(2/(k+1))
2)
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k/(k-1)
4.4
Solve for the vapor flow rate (assuming an orifice coefficient of 0.6 and a
Beta ratio of less than 0.2):
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2
k+1/k-1 0.5
2
k+1/k-1 0.5
2
k+1/k-1 0.5
W V = 2 (803) d {ρvkP1 [2/(k+1)]
W V = 2 (347) d {ρvkP1 [2/(k+1)]
W V = 2 (343) d {ρvkP1 [2/(k+1)]
}
4.5 (English)
}
4.5 (Metric)
}
4.5 (Metric)
Where:
WV
=
Vapor flow rate, lb/hr (kg/hr)
d
=
Tube I.D., inches
ρv
=
Vapor density at HP side operating conditions, lb/ft
3
(kg/m3)
k
=
Specific heat ratio for vapor = CP/CV
P1
=
High pressure side operating pressure, psia (bara,
kg/cm2A)
P2
=
Low pressure side accumulated relief pressure, psia
(bara, kg/cm2 A)
Pcf
=
Critical flow pressure, psia (bara, kg/cm A)
2
Vapor in Sub-Critical Flow (P2 greater than Pcf)
1)
Check for sub-critical flow by use of Equation 4.4
2)
Solve for the vapor flow rate
2
W V = 2 (1891) Yd C (∆PρV)
2
0.5
2
0.5
W V = 2 (816) Yd C (∆PρV)
W V = 2 (808) Yd C (∆PρV)
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0.5
4.6a (English)
4.6a (Metric)
4.6a (Metric)
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(See Crane Technical Paper 410, Equation 3.22)
Assume C = 0.6
Note: this is equivalent to a beta ratio less than about
0.2 and a Reynolds number above 1,000.
2
W V = 2 (1135) Yd (∆PρV)
0.5
2
0.5
2
0.5
W V = 2 (490) Yd (∆PρV)
W V = 2 (485) Yd (∆PρV)
4.6b (English)
4.6b (Metric)
4.6b (Metric)
Where:
Y is the expansion coefficient. An estimated value of Y can be
calculated with the following formula:
Y
=
[1 - 0.317(∆P) / P1] (Estimated Value)
4.7
or using the graphical solution on page A-21 in Crane Technical
Paper 410
C
=
Orifice coefficient (see Crane Technical Paper 410
page A-20)
Two Phase or Flashing Flow
A two-phase flow calculation method should be used in determining the flow rate
through a ruptured tube for flashing fluids or two phase fluids. Two phase flow
models developed by the AIChE Design Institute for Emergency Relief Systems
(DIERS) or others can be used for this purpose. However, the following approach
utilizes the methodology currently recommended in API RP-520 for sizing relief
valves for two-phase flow.
Calculate two-phase relief flow using the pure liquid and vapor equations (4.3
through 4.7). The value of P2 used for the liquid flow calculation should be the same
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as that used for the vapor flow equation; the greater of the critical flow pressure (Pcf)
or the backpressure (P2). Use the following steps:
1)
Check for two phase flow at tube rupture outlet conditions (Pbp is greater
than P2 )
Where:
Pbp = pressure at which vapor is first present, psia (bara, kg/cm2 A)
(Maximum = P1 )
2)
Use the k value for vapor at outlet conditions (P2 ) to calculate the critical
backpressure as
Pcf = [2/(k+1)]k/(k-1) ( Pbp )
4.7.2
4.8
3)
Calculate the vapor/liquid weight fractions at the greater of Pcf or P2.
4)
Calculate the relative areas required for the vapor and liquid weight
fractions.
5)
Prorate total flow based on an area of twice the ruptured tube cross
sectional area.
Steady State Relief Analysis
Relief loads for tube rupture cases are calculated based on the considerations
described in this section. A sample calculation is provided in Appendix B-3 and the
spreadsheet program for calculating flow through a tube rupture is provided in
Appendix D.
Two approaches are available for determining the size of a relief device, steady
state and dynamic analysis. If a steady state method is used, the relief device size
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should be based on the gas and/or liquid flow rate passing through the rupture, as
determined in Section 4.7.1. An exception occurs if the relief point is located at a
distance from the exchanger. For this special case, the equivalent volumetric flow
rate of the low pressure fluid based on the volume of high pressure fluid passing
through the rupture should be used if it results in a greater relief device area.
The configuration of the discharge piping and the contents (liquid or vapor) of the
low-pressure side of the exchanger should be considered in determining the
influence of piping in either eliminating the need for a relief device or in reducing
relieving requirements. Where the low-pressure side is in the vapor phase, full
credit can be taken for the vapor-handling capacity of the outlet and inlet lines,
provided that the inlet lines do not contain check valves or other equipment that
could prevent backflow. Where the low pressure side is liquid full, the effective
relieving capacity for which the piping system may be credited shall be based on the
volumetric flow rate of the low pressure side liquid that existed prior to the tube
rupture, unless a dynamic analysis is used.
Frequently, tube ruptures cases must be evaluated where high pressure flow is
being exchanged with cooling water. In that case, calculation of the credit for
displacement of cooling water can be based on analysis of the network flow. Due to
the large network, flow may be displaced through several exchangers. This can be
addressed as follows:
1)
Calculate the volume flow through the tube rupture, QTR.
2)
Locate the point in the unit header where normal cooling water volume flow,
QCW , exceeds QTR. If the header flow does not exceed QTR, then the
maximum possible credit is equal to the header cooling water volume flow
and a relief case exists.
3)
Assuming normal operating pressure at the location selected in Step 2,
calculate the pressure at the exchanger based on the properties of the fluid
flowing across the broken tube.
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4)
a)
If the calculated pressure is less than or equal to the exchanger
MAWP and QCW exceeds QTR, then there is no relief case.
b)
If the calculated pressure is less than or equal to the exchanger
MAWP x 1.1, a relief valve is required, but full credit can be taken for
QCW .
c)
If the calculated pressure exceeds MAWP x 1.1, then a location on
the cooling water system closer to the exchanger can be selected
and the calculation repeated to try to obtain a lesser credit.
The required volumetric relief load is QTR - QCW , where QCW is the largest
displaced flow for which credit can be taken.
In some cases the calculated volume flow through the tube rupture exceeds the
normal flow through the high pressure side of the exchanger. In those cases, an
assessment should be made to determine whether that high flow can be sustained
for a long enough time to be of concern. If there is a large volume of high pressure
vapor available to flow to the exchanger from vessels upstream and downstream of
the exchanger, then sustained flow at the high calculated rate should be considered.
Where it is unclear if sustained flow should be considered, then the client should be
advised and agree on the conclusion. Dynamic simulation can be used as a tool to
provide some guidance for this circumstance.
Tube rupture may also have the impact of eliminating heat transfer in the
exchanger. This may result in an added relief load due to loss of cooling which may
impact the total anticipated relief load.
4.7.3
Dynamic Relief Analysis
This type of analysis is recommended where there is a wide difference in design
pressure between the two exchanger sides. This is particularly recommended
where the low pressure side is liquid full and the high pressure side contains a gas
or fluid which will flash across the rupture. Modeling has shown that under these
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circumstances transient conditions may produce significant overpressure, even
when protected by a pressure relief device. This approach can sometimes reveal
alternative methods to relief device installations which can represent a preferred
solution. In addition, dynamic analysis can accurately account for acceleration
effects in the system.
Results from a dynamic simulation must be interpreted based on experience, with
consideration of the likelihood that a rupture probably does not occur
instantaneously. Therefore, the real impact of short term transients must be
carefully considered. Short term pressure spikes beyond exchanger hydrotest
pressure can often be tolerated without loss of containment.
4.7.4
Relief Devices and Locations
Where it is determined that a relief device is required for the particular installation,
rupture disks and/or pressure relief valves should be considered. While response
times of the devices should be considered, the relative concern of short term
transients should be weighed against the potential impacts on system operability or
vent disposal concerns that differ between relief device options. Rupture disks
installed for ruptured tube relief have sometimes prematurely leaked, resulting in
shutdowns and occasionally diversion of low pressure side flow to inappropriate
disposal locations.
It is preferable to locate a required relieving device as close as possible to the
exchanger, to avoid transient and steady state pressure buildup. Locating a relief
valve a significant distance from the exchanger, where the low pressure side is
liquid full, could lead to high transient pressure for a long enough period of time to
require sizing the relief valve for displaced liquid flow, as indicated in Section 4.7.2.
4.7.5
Double Pipe Exchangers
Units that use schedule pipe for the inner conduit or tube are no more likely to
rupture the inner pipe than any other pipe in the system; therefore, failure need not
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be considered a source of pressure relief requirement. Where gauge tubes are
used, it should be determined whether or not they are equivalent to schedule pipe.
4.8.
FIRE
In the absence of client guidelines, for all facilities in refineries and petrochemical plants,
the procedures given in API standards or recommended practices (RP 521 or API 2000)
are to be utilized, unless otherwise ruled by local (or country) regulations. The lead process
engineer should consult with the client to find and determine the suitability of any client
guidelines dictating the fire case relief load calculation methods. For overseas projects,
some countries use different calculation methods than API (for example, those countries
that have adopted Japanese standards). In the United States, some states (or cities)
dictate the use of methods in the NFPA codes. Therefore, it is important for the lead
process engineer to define the fire load calculation methods at the beginning of any new
project.
The following sections discuss the fire case relief load calculation procedures based on
API guidelines.
4.8.1
Basic Assumptions for Fire Case Relief Analysis
Protection of process plant systems from the impact of a fire involves a combination of
approaches which, for example may include the case of fireproofing, fire monitor
location, equipment location, emergency response by fire crews and equipment
depressurizing, in addition to the use of pressure relief devices. Determination of how
a system is impacted during a fire is difficult to quantify. However, through testing,
experience and theoretical calculations, various industrial, fire and safety organizations
have developed simplified approaches for sizing pressure relief systems to handle fire
case loads.
The following basic assumptions are used in calculating fire case relief loads:
o
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The process is assumed to be shut down and isolated from other vessels,
sources of process fluids or other relief paths.
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4.8.2
o
Liquid inventories are assumed to be at their normal control point, or if not
controlled, at their maximum normal operating level.
o
Heat input to process equipment is calculated based on empirical equations
which include parameters for adjusting heat flux to reflect special
circumstances associated with each installation.
o
All fire heat input is normally assumed to be available for vaporizing or heating
vessel contents, even though it may take a long time to heat up the vessel
contents.
Heat Flux Equations
4.8.2.1
Background
A discussion of the origins of heat flux equations in common use in the
United States is provided in API RP-521 and its associated references. In
the United States, fire loads in refinery applications are normally calculated
using an API equation provided in API RP-521 and by other equations for
calculation of relief loads in storage tanks.
The basic heat flux equation recommended by API RP-521 for equipment
containing liquid is:
Q = 21,000 F A0.82
Q = 155,426 F A0.82
0.82
Q = 37,140 F A
4.11 (English)
4.11 (Metric)
4.11 (Metric)
Where:
Q
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=
Total heat input to the wetted surface, Btu/h (kJ/h,
kcal/h)
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F
=
Environmental factor to account for vessel insulation
or special fire fighting capabilities (F = 1.0 for bare
vessel)
A
=
Vessel wetted area, ft2 (m2 )
21,000
=
Empirical coefficient from testing, including credits
normally encountered in refinery (155,426,37,140)
0.82
=
Empirical coefficient from testing
As indicated, the factor of 21,000 (155,426,37,140) includes credits for
favorable conditions normally encountered in process units. These conditions
are:
o
Sloping of the grading and drainage systems so that flammable liquid
will not pool under process vessels.
o
Fire fighting activity is expected to begin soon after a fire starts.
If good drainage and fire fighting facilities are not present, then the coefficient
2
2
2
of 21,000 Btu/h-ft (155,426 kJ/h-m , 37,140 kcal/h-m ) should be increased to
2
2
2
34,500 Btu/h-ft (255,277 kJ/h-m , 61,000 kcal/h-m ). This coefficient appears
as the required fire heat input coefficient in some standards related to LPG
storage vessels (NFPA 58, OSHA 1910.110), which do not recognize drainage
credits.
The F factor in the API equation is primarily used to account for insulation
credits. No credit is allowed for water spray systems.
The F factor is calculated as:
F = k (1660° - Tf)
21000 t
F = k (904° - Tf)
2389 t
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4.12 (English)
4.12 (Metric)
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F = k (904° - Tf)
570 t
4.12 (Metric)
Where:
k
=
Insulation thermal conductivity calculated at the mean
temperature of Tf and 1660 °F (904°C), Btu in/ft h °F (kJ
cm/ m2 h °C, kcal cm/m2 h °C).
(k can be estimated for mineral wool and calcium silicate
insulation as: k = .03 + 2 x 10-4 Tf, kcal cm/m2 h °C)
2
Tf
=
Fluid temperature, °F (°C)
t
=
Insulation thickness, in (cm).
Fire case heat absorption calculations for storage tanks are governed by
criteria provided by API, NFPA and OSHA in the United States. These
methodologies differ primarily in the credits allowed for environmental factors
and in the extent that increasing tank size is considered to reduce heat flux
over the entire wetted surface.
4.8.3
Determination of Relief Loads for Equipment Containing Liquid
Fire relief loads for equipment containing liquid are calculated using the procedures in
this section. A sample calculation is provided in Appendix B-4 and the spreadsheet
program for this calculation is provided in Appendix D.
4.8.3.1
Determination of Wetted Area
o
Vessels and Storage Tanks
Wetted area will be calculated based on the following guidelines.
Generally, only the wetted area up to a height of 25 feet (7.6 meters)
above the fire surface will be considered, based on API RP 521 criteria.
However, for horizontal vessels, any wetted area above 25 feet (7.6
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meters) will also be included, up to the maximum horizontal diameter
(½ full). For horizontal drums 6 feet (1.9 meters) or less in diameter,
the entire area will be considered wetted. Include the bottom head of
vertical vessels in the wetted area, even if the skirt is fireproofed.
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Equipment
Wetted Area
Liquid full vessels
Total vessel surface
Accumulators/
Separators (w/liquid
level control)
Base wetted surface on normal
operation level. If unknown
assume 1/2 full.
Surge drums
Base wetted surface on maximum
operation level shown on P&ID. If
unknown, assume maximum operation
level at 75% of diameter for horizontal
drums and 75% of T/T height for vertical
drums.
Knockout drums
(no normal level)
Use liquid level of 25% of diameter for
horizontal drums and 25% of T/T height
for vertical drums.
Columns
Use normal liquid level in tower plus
assume 4 in (100 mm) liquid height per
tray (based on total tower cross section)
drains into tower bottom to obtain total
liquid height in column.
Atmospheric Storage tanks
Per API 2000, as follows:
Spheres - Greater of 55 percent of total
area or total area up to 30 feet (9.1
meters).
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Horizontal tank - Greater of 75 percent of
total area or total area up to 30 feet (9.1
meters).
Vertical tank - Total area up to 30 feet
(9.1 meters).
Note: Refer to API 2000 for heat input
equations, which vary with tank area and
design pressure.
Spheres
o
Wetted surface up to greater of
maximum horizontal diameter or to the
height of 25 feet (7.6 meters).
Shell and Tube Exchangers
No specific guidelines are provided for heat exchanger wetted areas.
Wetted area should be calculated based on the outside area of the
heat exchanger adjusted for the normal volume fraction of liquid in the
shell side stream. The calculation would then be similar to the
calculation for drums and vessels. Consideration should also be given
to the potential to drain liquid to or from the exchanger after the system
is shut down.
If it is possible to transfer heat from the shell side of the exchanger to
the tube side fluid under relieving conditions, then it should be
considered that the fire heat flux could go either to the shell side or
tube side liquid. Fire case heat flux into tube side liquid should be
considered if the bubble point of the tube side fluid at tube side MAWP
is less than the bubble point of the shell side fluid at accumulated relief
pressure.
o
Air Cooler
Air coolers will be excluded from fire case relief load calculations.
Although API RP 521 discusses wetted areas that could be used to
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estimate fire case relief loads for air coolers, in practice calculated
loads using such an approach are extremely high due to the large
surface area provided by the exposed tubes. Design of a relief valve to
accommodate very large, short term relief loads is frequently
impractical and generally not desirable. Protection of air coolers
against fire impacts is preferably approached using other techniques.
Most important is the initial design of the air cooler installation. Free
drainage of the exchanger should be designed in to the maximum
extent practical. Location of the air cooler above and away from
regions of likely fire exposure should be considered. If design factors
alone are not considered sufficient to mitigate concerns with specific air
coolers, then additional fire fighting protection, such as the used of
water deluge systems, can be considered.
o
Miscellaneous Equipment
Other miscellaneous equipment such as filters which are ASME code
stamped should also be provided with relief protection from fire. The
wetted area for such equipment is calculated similarly to that for
vessels.
o
Piping
Piping around equipment is not normally included in the wetted area
used to calculate the fire case relief load. However, in some cases,
liquid-full piping around equipment represents a significant portion of
exposed surface area and should be considered in calculating heat flux
during a fire. In the case where significant insulation credit is taken for
the equipment and the associated piping remains uninsulated, a
nominal 25 feet (8 meters) of the single largest inlet line and the single
largest outlet line which are likely to contain liquid shall be included in
the wetted area. In new systems, piping insulation suitable for fire
protection should be provided around any equipment for which fire
protection insulation is being provided.
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4.8.3.2
Insulation Credit
While the API and HPGSA equations provide a means of calculating the
credit which can be taken for insulation during a fire, the installed insulation
system must meet certain standards as indicated in API RP-521 before any
credit is taken. Furthermore, practical limits on the extent of insulation credit
taken are appropriate to avoid establishing fire case relief loads which could
be exceeded due to normal wear and tear of equipment insulation.
In order to take credit for insulation properties to reduce fire case relief loads,
the installation must withstand both the temperature of the fire and the
mechanical forces of fire water. As a minimum, this requires either the use
of stainless steel wire to bind the insulation underneath the sheet metal
sheathing or the use of fire-resistant banding to hold the sheet metal to the
vessel. For new installations, fire-resistant sheet metal should be used in
combination with stainless steel wiring and/or banding. Both stainless steel
and galvanized carbon steel sheet metal meet the requirement for fire
resistant jacketing. While galvanized carbon steel is easier to work with and
less expensive than stainless steel, the use of galvanized steel to jacket
stainless steel equipment is usually avoided.
In order to take credit for insulation, the insulation must withstand fire case
temperatures. Foam, plastic and fiberglass type insulations frequently used
for cold services are unsuitable fire protection purposes.
Typical thermal conductivities for suitable insulation materials result in an
order of magnitude reduction in calculated fire case heat flux from that
calculated for bare equipment with minimum normal installation thickness.
Caution should be used to avoid taking excessive credit which can result in a
design without sufficient conservatism.
In order to effectively manage and control fire protection insulation, its use
must be properly documented and placed under administrative control. To
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accomplish this, fire-protection insulation must be clearly identified on P&ID's
and in pressure relief device engineering files.
4.8.3.3
Liquid Filled Systems
The characteristics of liquid filled systems must be considered in evaluating
fire pressure relief systems. Initially there is no vapor space and before
boiling begins, hydraulic expansion occurs and liquid will release. Once
boiling begins, there is little or no vapor/liquid space and the relief stream
may be vapor, liquid or a two-phase stream. Where liquid filled equipment is
directly connected to a vessel which contains a vapor space, provided there
is a clear flow path, often the relief valve on the pressure vessel with the
vapor space can be used to protect the liquid filled equipment.
If the pressure relief valve is located where vapor produced by boiling cannot
immediately reach the valve, it must be sized for a liquid flow equal in volume
to the generated vapor which is displacing it. This can occur because the
relief valve is not mounted at the high point of the vessel being protected or
because a single relief valve may be protecting several pieces of liquid filled
equipment.
While it is clear that two-phase flow can result during relief of a liquid full (or
even partially full) vessel during a fire, relief valves will be sized based on
single phase flow into the relief valve (i.e. liquid expansion, liquid
displacement, or vapor flow) for fire relief unless there is other evidence
which suggests that this approach is not conservative for the particular
application. However, provisions should be considered for handling of about
1/3 of the vessel contents as liquid passing into the valve for purpose of
establishing the disposal system design. Further guidance in estimating the
total quantity of liquid carryover is provided as part of DIERS (Design
Institute for Emergency Relief Systems) technology. Depending on the
shape of the vessel or equipment item, the rate of vapor generation, the
nature of the fluid, and the fill level, carryover may occur while the vessel is
between 60% to 95% full.
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4.8.3.4
Protection of System with Individual Relief Valve
It is possible to protect multiple pieces of equipment with one relief device for
fire case conditions, provided certain conditions are met. Boundary limits for
protection of a system are first established based on points of isolation which
can occur during a fire. These include check valves, pumps, control valves,
and battery limit block valves. Each of these represent potential blockage
points which should be considered as isolating section(s) of the process
system during a fire.
Within that section of the process, it is possible for one relief device to
provide fire protection for multiple equipment items if the relief device is sized
to handle the relief loads, a path is always open from each item of equipment
to the relieving device, and the effect of liquid heat and flowing pressure drop
is properly considered in determining relief valve set pressure (so that all
items of equipment are properly protected).
If multiple relief devices are used within a single, isolated section of the
process, and they are set at significantly different relief pressures, a check
must be made to identify where the fire case relief loads will relieve.
Sometimes the entire load must be relieved from a lower set pressure relief
valve.
4.8.3.5
Maintenance Isolation
It is generally recommended to minimize the number of maintenance
isolation valves provided for equipment. However, if it is found necessary to
provide maintenance isolation valves, consideration must be given to
providing fire relief protection for equipment which can be isolated.
Vessels which can be isolated are normally protected by a dedicated relief
valve. Frequently, however, administrative procedures are used to avoid
providing separate relief devices for individual exchangers, and other small
equipment items subject to isolation. The basis of these procedures is to
require immediate drainage and venting of the isolated equipment. As with
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other administrative procedures, the practice to be followed must be carefully
documented and included in training procedures. All equipment items to be
handled in this manner must be clearly identified. Application of this
procedure to active spare equipment (including certain exchangers and
filters) must be carefully reviewed with operations to obtain agreement that
the equipment will remain free of liquids until placed into service, since
frequently stand-by equipment is maintained liquid full.
4.8.3.6
Determination of Latent Heat for Boiling Applications
For fluids boiling at relief pressure at moderate temperatures, well below the
critical point, the fire case relief load is calculated as the heat flux to the fluid
divided by the latent heat of the fluid. Latent heat has a straightforward
single value at relief pressure for a pure component. However, refinery
processes generally involve processing of multiple components. The
following rules are adopted for the liquid with multiple components.
•
•
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For a given liquid in a system other than fractionation tower, latent heat is
calculated based on the point of 5 mol % vaporization, as the difference
in enthalpy between the vapor stream and liquid steam expressed in
Btu/lb (kJ/kg, kcal/kg). While this is not a unique or universally
standardized methodology, it provides a reasonable approach for the
following reasons:
-
Avoids excessive impact of minor amounts of light ends which may
flash off at the initial boiling point.
-
Generally, for the same heat input, relief of the lighter components
results in the largest calculated relief device.
An exception to this approach is made for tower systems, where the
same latent heat used for other tower loads, such as power failure, can
be used to simplify calculations. Adjustment for the different
overpressure is not required because the latent heat change is
considered negligible. The latent heat is determined at the dew point
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temperature for the reflux stream. This represents light ends
vaporization at the beginning of a fire and is conservative as discussed
above.
•
4.8.3.7
In the absence of simulation data, it is also possible to estimate relief
properties from Figure A-3 in API RP-521 (Third Edition). While the data
provided there is for pure paraffin compounds, conservative estimates of
latent heat can be obtained from the chart. The wider the range of
hydrocarbons in the mixture, the less accurate the chart estimates will
be.
High Boiling Point Fluids
High boiling point fluids may not vaporize at relieving pressure before
reaching extreme temperatures. For practical purposes, a pressure relief
device does not add protection to such a system at the boiling point
temperature. As a working definition, this point is reached once a
temperature about 230 °F (110 °C) beyond the point where the equipment
material stress curve begins to decline is reached. For carbon steel this is
about 850 °F (455 °C) and for chrome-moly steel this is about 1000 °F (540
°C). Beyond these temperatures, the following approach should be used:
4.8.3.8
o
Size the relief valve based on liquid expansion if the equipment is liquid
full.
o
Size the relief valve based on vapor expansion if the equipment is not
liquid full.
o
If thermal cracking is known to occur and can be quantified (or at least
estimated), add that load to the loads calculated by expansion.
Latent Heat of Hydrocarbon/Water Mixtures
Free water is sometimes present in addition to the hydrocarbon liquid phase.
For equipment containing substantial quantities of water such as wash
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drums, treater, coalescers, and desalters, a mixed phase equilibrium
calculation is used to calculate the relief load. This calculation is based on
the principle that each fluid contributes its vapor pressure independently.
The presence of water can significantly reduce the calculated relief load
compared to pure hydrocarbon due to the high latent heat of water.
This method should not be used for vessels containing minor amounts of
water or where free water is only occasionally present since a nonconservative load will be predicted. This includes vessels where the water
phase is normally contained entirely in a boot. For these cases vaporization
of the hydrocarbon phase only should be considered.
4.8.3.9
Critical or Super-Critical Fluids
Fluids which may reach or exceed their critical point during the pressure rise
to relieving conditions in a fire require special consideration in calculating the
relief loads. Fluids at or above critical pressure have no latent heat and
behave somewhat like vapors.
Detailed calculation of the relief load under fire conditions for critical fluids is
based on expansion of the contents in the equipment as follows:
o
Calculate heat input to the vessel assuming the entire volume is full of
a "super-critical liquid."
o
Calculate the heat input using the standard equation for a short time
interval and calculate the required vessel fluid mass relieved to
maintain vessel pressure.
o
Perform the calculation for several intervals until it is clear that a
maximum relief rate has been found.
For hydrocarbon liquids, a simplified (conservative) approach is to use a
minimum "equivalent" latent heat of 20 Btu/lb (46 kJ/kg, 11 kcal/kg) to
calculate the relief load. This method can be used to calculate relief rates for
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new relief valves in hydrocarbon systems where the fluids are near critical or
are super-critical to reduce calculation time, unless detailed calculations
appear warranted. It should be considered that this “equivalent” latent heat
is intended to be conservative, and can be much greater for many
hydrocarbons and for conditions somewhat removed from critical.
4.8.4
Relief Loads for Vessels Containing Vapor
Vapor filled vessels will not absorb heat in the same manner that vessels containing
liquid will. Internal pressure will rise in proportion to the absolute vapor temperature in
the vessel. Low heat transfer rate between the vessel wall and the vapor in the vessel
can cause the vessel metal temperature to rise to very high levels. As a result, the
metal may have insufficient strength to contain the internal pressure. Therefore, the
use of a pressure relief device alone may not be adequate to protect the vessel,
requiring depressuring and or special fire-fighting provisions.
The basic method for sizing pressure relief systems for vapor filled vessels exposed to
fire is specified in API RP-520 as:
0.5
W = 0.1406 (MP1)
1.25
(Tw - T1)
A
4.13 (English)
T1 1.1506
W = 2.772 (MP1)0.5 (Tw - T1)1.25 A
4.13 (Metric)
T1 1.1506
W = 2.745 (MP1)0.5 (Tw - T1)1.25 A
4.13 (Metric)
T1 1.1506
Where:
Chap4-r1.doc
W
=
Relief flow, lb/h (kg/h)
A
=
Exposed surface of the vessel, ft (m )
P1
=
Relieving pressure, psia (bara, kg/cm2 A)
2
2
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M
=
Molecular weight of vapor
Tw
=
Vessel wall temperature, °R (°K)
[recommended as 1560 °R (865°K) for carbon steel; adjust for
other materials to reflect recommended maximum temperature]
T1
=
Relieving temperature, determined by the following operation:
T1 = Tn P1
4.14
Pn
Where:
Pn = Normal operating pressure, psia (bara, kg/cm2 A)
Tn =
Normal operating temperature, °R (°K)
This equation often predicts very conservative relief loads, because the vessel is
assumed to be uninsulated, vapor is assumed to be an ideal gas and the temperature
of the vapor in the vessels is assumed to be uniform. A more detailed method is
referenced in API RP 520, which considers the dynamic effects of heat transfer to the
vessel and allows the effect of insulation to be accounted for. However, this method
should not normally be used, since the simplified method rarely results in the
calculation of major relief loads or very large relief valve sizes. The exposed area, A,
is calculated using the same fire exposure height limit of 25 feet (7.6 meters) used for
wetted vessels in Section 4.8.3.1.
4.8.5
Depressuring
Controlled depressuring of the vessel reduces internal pressure and stress in the
vessel walls. It also guards against the potential addition of fuel to the fire should
the vessel rupture. The design of depressuring systems should recognize the
following factors:
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o
Manual controls near the vessel may be inaccessible during an emergency.
o
Unless anticipated, automatic controls could fail in a direction that would
prevent depressuring (for example valves that fail closed).
o
Early initiation of depressuring is desirable to limit vessel stress to
acceptable levels commensurate with the vessel wall temperature that
results from a fire.
o
Safe disposal of vented streams must be provided.
o
No credit is recommended when safety valves are being sized for fire
exposure.
Further information on depressuring is provided in API RP 521 (Third Edition),
paragraph 5.2.3
Generally an external fire does not cause rupture of a vessel below the liquid level
because the vessel shell temperature is limited by good heat transfer to the boiling
liquid. However, for vessel portions above the liquid level, for gas filled vessels, and
for internally insulated or coked vessels, an external fire may cause shell
overheating and rupture in a relatively short time even with a properly functioning
pressure relief valve system. Therefore, the need for additional precautions should
be reviewed with the client at the beginning of any new project. These precautions
may include one or more of the following options:
o
Provide fire protection facilities.
o
Consider the need for enhanced fire protection such as water sprays or
water deluge systems installed per API Publication 2030 or NFPA 15. No
credit is to be taken on pressure relief valve sizing for water application
facilities.
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o
Provide a vessel emergency depressuring system to enable depressuring of
the vessel to 50% of the relief valve set pressure or 100 psi (6.9 bar, 7
2
kg/cm ), whichever is less, in 15 minutes. The emergency depressuring rate
should be determined as outlined in API RP 521 (Third Edition), Paragraph
3.16. Do not take a credit for this quantity in sizing the required relief valve.
However, a credit may be taken for total relief to the relief header.
o
Provide adequate surface drainage away from vessels and equipment.
o
Consider reducing the fire heat input rate and the vessel shell temperature
increase rate by using vessel insulation.
o
Consider reducing the fire heat input rate by increasing the vessel elevation.
To reduce the internal pressure in equipment involved in a fire, vapor must be
removed at rate which compensates for the following occurrences:
o
Vapor generated from liquid by heat input from the fire.
o
Density change of the internal vapor during pressure reduction.
o
Liquid flash due to pressure reduction. (This factor applies only when a
system contains liquid at or near its saturation temperature).
The total vapor load for a system to be depressured may be expressed as the sum
of the individual occurrences for all equipment involved as follows:
=
wf + wd + wv
wt
=
Total vapor to be vented, lb (kg)
wf
=
Vapor generated by heat input from the fire, lb (kg)
wt
4.15
Where:
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wd
=
Vapor produced from expansion to a lower pressure, lb (kg)
wv
=
Liquid flash due to pressure reduction, lb (kg)
Vapor generation due to the fire is calculated in accordance with the procedures in
the preceding paragraphs.
The following formula is used to calculate vapor produced from expansion to lower
pressures:
wd = V[(P1M1/RZ1T1) - (P2M2/RZ2T2)] = V[ρ1 - ρ2]
4.16
Note: Subscript 1 designates initial conditions and subscript 2 designates
final conditions in systems being depressured.
Where:
V
=
System volume, ft3 (m3 )
R
=
Ideal gas constant
P
=
Fluid pressure, psi (bar, kg/cm )
M
=
Molecular weight
Z
=
Compressibility Factor
T
=
Fluid temperature, °R (°K)
ρ
=
3
3
Fluid density, lb/ft (kg/m )
2
For pure components or narrow boiling range mixtures, the amount of liquid flash
may be conservatively estimated by equating the heat of the flashed vapor with the
heat loss of the average liquid quantity. The simplified formula is as follows:
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wv = (wi - 0.5wf)(2Cp∆tb)/(2λ + Cp∆tb)
4.17
Where:
wv
=
Quantity of vapor to be removed due to flashing liquid, lb (kg)
wi
=
Quantity of liquid initially in system, lb (kg)
wf
=
Quantity of vapor to be removed due to heat from the fire, lb (kg)
Cp
=
Specific heat of the liquid, Btu/lb °F (kJ/kg °C, kcal/kg °C)
∆tb
=
Bubble point temperature change, °F (°C)
λ
=
Latent heat of vaporization, Btu/lb (kJ/kg, kcal/kg)
For wide boiling range mixtures, a series of simplified adiabatic flash calculations
must be made between initial pressure and final pressure while neglecting the
simultaneous fire effect. The simplified adiabatic flash calculation is a stepwise
procedure which yields a weight fraction flashed from the liquid quantity originally in
the system. It is assumed that vapors flashed in each step are totally removed from
the system to be depressured before the next step occurs. The correction for the
fire is made after completion of the simplified flash calculation.
o
To determine the approximate amount of liquid vaporized from a mixture, an
equilibrium phase diagram is required and a graphical solution is employed.
The procedure uses the following equation:
∆t = λ∆wv/(wi - Cp∆wv) for each flash step
o
Chap4-r1.doc
4.18
The weight of liquid flashed due to depressuring during a simultaneous fire is
then estimated by the following equation:
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wv = (wi - 0.5wf)(X)
4.19
Where:
X
=
Weight fraction of liquid flashed
The following criteria are used for setting pressure limits and time duration for
depressuring:
o
The initial pressure is usually the pressure relief valve set pressure.
o
The final pressure is taken as 100 psig (6.9 barg, 7.0 kg/cm2 G) or 50% of
the relief valve set pressure, whichever is less.
o
The depressuring time to be taken is 15 minutes. If the arithmetic average
flow is used to calculate depressuring time, the calculated depressuring time
is divided by two-thirds to compensate for exponential characteristics of flow.
The suggested method for sizing depressuring system lines was developed by C.
E. Lapple. It employs a theoretical mass flow based on ideal nozzle and isothermal
flow conditions and the following formulas used with Figure 4.1 (“Isothermal Flow of
Compressible Fluids through Pipes at High Pressure Drops”):
Gc = 12.6 P1 (M/T)0.5
Gc = 665 P1 (M/T)
0.5
Gc = 652 P1 (M/T)
4.20 (English)
4.20 (Metric)
4.20 (Metric)
NT = fL/Di + K
4.21
0.5
Where:
Gc
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=
2
2
Critical Mass Velocity, lb/ft sec (kg/m sec)
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4.9
G
=
2
2
Mass velocity, lb/ft sec (kg/m sec)
M
=
Molecular weight
T
=
Fluid Temperature, °R (°K)
f
=
Moody or Darcy Friction Factor, no units
L
=
Length of pipe, ft (m)
Di
=
Inside diameter of the pipe, ft (m)
K
=
Velocity head or resistance coefficient, no units
P1
=
Source pressure, psia (bara, kg/cm A)
P2
=
Destination Pressure, psia (bara, kg/cm2 A)
2
CHEMICAL REACTIONS
Calculating a relieving rate for a chemical reaction requires considerable knowledge of
reaction rates and reaction kinetics. If a chemical reaction can approach the conditions of
an explosion, the consideration in API RP 521 for internal explosions should be applied.
The particularly hazardous elements in exothermic reaction are the acceleration of the
reaction rate with rising temperature, which rapidly produces extremely high rates of energy
release, and the release in many instances of large volumes of non-condensables once the
temperature rises to excessive levels and decomposition reactions begin to occur. Under
these circumstances, normal overpressure relief may be insufficient and of little value; rapid
relief before pressure and temperature rise to exponentially accelerating levels is a
requirement in coping with possible runaway reactions.
If it were possible to devise a fully reliable system, a good method to prevent destructive
overpressures in such reactions would be to add enough volatile fluid to absorb excess
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reaction heat and cool the reacting material to safe levels. This concept of internal
refrigeration obviously requires a large, but calculable, relieving capacity. It represents an
instance of using a volatile admixture as a deliberate strategy to protect against more
serious overpressure situations. If the operating conditions result in sufficient cooling from
depressuring, vapor depressuring is an even more obvious means of averting situations
that involve runaway reactions and destructive overpressure.
4.10
ATMOSPHERIC STORAGE TANK PROTECTION
API 2000, 4th Edition, Sept. 1992 covers the normal and emergency venting requirements
for above ground liquid petroleum storage tanks and aboveground and underground
refrigerated storage tanks designed for operation at pressure from vacuum through 15 psig
2
(1.03 barg, 1.05 kg/cm G).
4.10.1 Relief Device Accumulation
4.10.1.1 Pressure
Under normal conditions, pressure-relieving devices must have sufficient
flow capacity to prevent the pressure from rising more than 10 percent
above the maximum allowable working pressure. Under fire emergency
conditions, the devices shall be capable of preventing the pressure from
rising more than 20 percent above the maximum allowable working
pressure. Pallet valves typically require a set pressure at ½ MAWP to
allow for accumulation.
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4.10.1.2 Vacuum
The set and relieving pressures for vacuum relief are established to
prevent damage to a tank and must limit vacuum to a level no greater than
that for which a tank has been designed. The vacuum-relieving devices of
a tank shall be set to open at a pressure or partial vacuum that will ensure
that the partial vacuum in the tank will not exceed the partial vacuum for
which the tank has been designed when the inflow of air through the
devices is at its maximum specified rate. For pallet operated valves the set
pressure is typically 1/20th of the maximum vacuum gauge.
4.10.2 Nonrefrigerated Aboveground Tanks
4.10.2.1 Causes of overpressure
When the possible causes of overpressure or vacuum in a tank are being
determined, the following circumstances must be considered:
o
Liquid movement into or out of the tank
Inbreathing will result from the outflow of liquid from a tank.
Outbreathing will result from the inflow of liquid into a tank and from
the vaporization, including flashing, of the feed liquid, that will occur
because of the inflow of the liquid. Flashing of the feed liquid can be
significant for feed that is near or above its boiling point at the
pressure in the tank. A frequent cause of damage to tankage occurs
where equalization flows between connected tanks are in excess of
the design process inflows and outflows. An example of this occurs
when two or more tanks have a common suction line to a transfer
pump. If two tanks are left in communication during switching of
services, the liquid in the tank with the higher level will flow into the
tank with the lower level. These rates can be very high and must be
considered where this equalization can occur.
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o
Tank breathing due to normal atmospheric pressure and temperature
changes
Inbreathing will result from the contraction or condensation of vapors
that is caused by a decrease in atmospheric temperature.
Outbreathing will result from the expansion and vaporization that are
caused by an increase in atmospheric temperature (thermal
breathing).
o
Fire exposure
Outbreathing will result from the expansion of the vapors and
evaporation of the liquid that occur when a fire is in the vicinity of a
tank.
o
Other circumstances resulting from equipment failure and operating
errors
When the possible causes of overpressure or vacuum in a tank are
being determined, other circumstances resulting from equipment
failures and operating errors must be considered. The following
circumstances, along with any other possible controlling factors,
should be considered by the designer:
•
Pressure Transfer Blowoff
Where liquid transfer into a tank is accomplished by pushing
with a gas pad on the source, a surge of the pad gas is
normal at the end of transfer. The receiving tank must
account for this source of potential overpressure with
adequate venting capacity.
•
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Inert Pads and Purges on Tankage
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Inert pads and purges are sometimes provided on tanks to
prevent contamination, maintain nonflammable vapor spaces
in the tank and suppress vapor emissions from the tanks. A
typical pad or purge system contains a supply regulator and a
back pressure regulator. Failure of either of these regulators
needs to be accounted for in the tank venting requirements.
•
External Heat Transfer Devices
Steam, tempered water, and hot oil are common heating
media for tanks whose contents must be maintained at
elevated temperatures. Failure in the open position of the
heat media control valve, sensing device or control system
must be accounted for in the tank venting requirements.
•
Internal Heat Transfer Device
Mechanical failure of a tank’s internal heating or cooling
device can expose the contents of the tank to the heating or
cooling medium used in the device. In low-pressure tanks, it
can be assumed that the failure flow direction will be into the
tank. Chemical compatibility of the tank contents and the heat
transfer medium must be considered. Relief of the heat
transfer medium (for example, steam) may be necessary.
The disposition of the tank contents until the device can be
repaired or replaced must also be considered.
•
Vent Treatment Systems
If a vapor from a tank is collected for treatment or disposal by
a vent treatment system, the vent collection system may fail.
This failure must be evaluated. Failures affecting the safety
of a tank can include back pressure developed from problems
in the piping (liquid-filled pockets and solids buildup), other
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equipment relieving into the header, or blockage due to
equipment failure. An emergency vent to atmosphere, set at
a higher pressure than the vent treatment system is normally
used. For toxic or hazardous vapors, a fail-safe vent
treatment system should be considered.
•
Utility Failure
Local and plant-wide power and utility failures must be
considered as possible causes of overpressure or vacuum.
Loss of electrical power will directly affect any motorized
valves or controllers and may also shut down the instrument
air supply. Also, cooling and heating fluids may be lost during
an electrical failure. See Section 4.1.3.
•
Change in Temperature of the Input Stream to a Tank
A change in the temperature of the input stream to a tank
brought about by a loss of cooling or an increase in heat input
may cause overpressure in the tank and must be considered
in the vent sizing.
•
Chemical Reactions
Chemical reactions can provide significant heat input resulting
in relief requirements. See section 4.2.8 and 4.9.0 for further
discussion.
•
Liquid Overfill Protection
Relief due to liquid overfilling can be mitigated with proper
instrumentation. See API RP 620 and API RP 2350 for
further information.
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•
Atmospheric Pressure Changes
A rise or drop in barometric pressure is a possible cause of
vacuum or overpressure in a tank.
•
Control Valve Failure
Failure of a control valve on the liquid line to a tank must be
considered because a control valve failure may adversely
affect the flow of material to a tank. Increased flow may
cause hotter fluid to enter the tank if a cooler is in the feed
line. Also, consider if failure of a control valve in the feed line
could cause loss of a liquid seal to the supply source resulting
in vapor flow directly to the tank.
4.10.2.2
Determination of Required Venting Rates
o
Normal Venting Requirements
Normal venting requirements (inbreathing/vacuum relief and
outbreathing/pressure relief) must satisfy the maximum
requirements for liquid flows into and out of the tank as well as
thermal breathing caused by changes in atmospheric temperature.
Credit may be taken for vapor generation or condensation where it
provides part of the venting requirements. API 2000, 4th Edition,
paragraph 2.4.2, offers some specifics for liquids with flash points
above and below 100 °C (38 °C). Consider the effect of any
noncondensables present.
Note that the thermal inbreathing requirements given by Table 2 of
API 2000 are approximately equivalent to a rate of change in
ambient temperature of 68 °F (38 °C) per hour. While this may
seem excessive, it reflects a change of about 18 °F (10 °C) in 15
minutes, which is not uncommon as storm fronts move through. It
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also considers the impact of sudden cold rainfall on the shell of the
tanks.
o
Emergency Venting Requirements
Emergency venting is required to satisfy the need for protection
when tanks are exposed to fire. The required emergency relief rate
can be provided by designing the tank to tear open at a frangible
roof-to-shell attachment. The design requirements for this weak
link are covered in API 650.
4.10.3 Refrigerated Aboveground and Belowground Tanks
4.10.3.1
Causes of Overpressure or Vacuum
When the possible causes of overpressure or vacuum in a refrigerated
tank are being determined, the following circumstances must be
considered:
o
Liquid movement into or out of the tank
Inbreathing will result from the outflow of liquid from a tank.
Outbreathing will result from the inflow of liquid into a tank and from
the vaporization, including flashing, of the feed liquid, that will occur
because of the inflow of the liquid. Flashing of the feed liquid can be
significant for feed that is near or above its boiling point at the
pressure in the tank. Vapors generated during the filling operation
also may come from a warm fill line, heat leak and fill pump work,
cooldown of the tank and fill line, and displacement by the incoming
liquid.
o
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Vacuum can develop in a tank when the ambient temperature drops
and causes a reduction in the temperature and vapor pressure of the
liquid in the tank.
o
Fire exposure
Outbreathing will result from the expansion of the vapors and
evaporation of the liquid that occur when a fire is in the vicinity of a
tank.
o
Other circumstances resulting from equipment failure and operating
errors
When the possible causes of overpressure or vacuum in a tank are
being determined, other circumstances resulting from equipment
failures and operating errors must be considered. The following
circumstances, along with any other possible circumstances, should
be considered by the designer:
•
Pressure Transfer Blowoff
Where liquid transfer into a tank is accomplished by pushing
with a gas pad on the source, a surge of the pad gas is
normal at the end of transfer. The receiving tank must
account for this source of potential overpressure with
adequate venting capacity.
•
Inert Pads and Purges on Tankage
Inert pads and purges are sometimes provided on tanks to
prevent contamination, maintain nonflammable vapor spaces
in the tank and suppress vapor emissions from the tanks. A
typical pad or purge system contains a supply regulator and a
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back pressure regulator. Failure of either of these regulators
needs to be accounted for in the tank venting requirements.
•
Heat Transfer Devices
For a tank with a cooling jacket or coils, the loss of flow of the
coolant must be considered.
•
Internal Heat Transfer Device
Mechanical failure of a tank’s internal heating or cooling
device can expose the contents of the tank to the heating or
cooling medium used in the device. In low-pressure tanks, it
can be assumed that the failure flow direction will be into the
tank. Chemical compatibility of the tank contents and the heat
transfer medium must be considered. Relief of the heat
transfer medium (for example, steam) may be necessary.
The disposition of the tank contents until the device can be
repaired or replaced must also be considered.
•
Vent Treatment Systems
If a vapor from a tank is collected for treatment or disposal by
a vent treatment system, the vent collection system may fail.
This failure must be evaluated. Failures affecting the safety
of a tank can include back pressure developed from problems
in the piping (liquid-filled pockets and solids buildup), other
equipment relieving into the header, or blockage due to
equipment failure. An emergency vent to atmosphere, set at
a higher pressure than the vent treatment system is normally
used. For toxic or hazardous vapors, a fail-safe vent
treatment system should be considered.
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•
Utility Failure
Local and plant-wide power and utility failures must be
considered as possible causes of overpressure or vacuum.
Loss of electrical power will directly affect any motorized
valves or controllers and may also shut down the instrument
air supply. Also, cooling fluids may be lost during an electrical
failure.
•
Change in Temperature of the Input Stream to a Tank
A change in the temperature of the input stream to a tank
brought about by a loss of cooling or an increase in heat input
may cause overpressure in the tank and must be considered
in the vent sizing. A reduction in vapor pressure brought
about by the introduction of subcooled product into the vapor
space may create vacuum conditions. Relief for “rollover”
conditions is not provided. Proper design and operation are
expected to prevent this condition.
•
Heat Inleak
Heat inleak to a refrigerated tank can cause overpressure in
the tank.
•
Liquid Overfill Protection
The use of redundant level instrument can provide additional
protection against liquid overfilling. See API RP 620, API RP
2350, and API 2510 for further information
•
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A rise or drop in barometric pressure is a possible cause of
vacuum or overpressure in a tank.
•
Control Valve Failure
Failure of a control valve on the liquid line to a tank must be
considered because a control valve failure may adversely
affect the flow of material to a tank. Increased flow may
cause hotter fluid to enter the tank if a cooler is in the feed
line. Also, consider if failure of a control valve in the feed line
could cause loss of a liquid seal to the supply source resulting
in vapor directly to the tank.
•
Pump Recycle
Vapor generated during the operation of a pump on recycle or
during recirculation can cause tank overpressure.
4.10.3.2
Determination of Required Venting Rates
Venting requirements are set forth for the following conditions:
o
Inbreathing resulting from maximum outflow of liquid from the tank.
Pressure relief devices shall be suitable to relieve the flow capacity
determined for but not limited by the largest single contingency or
any reasonable and probable combination of contingencies,
assuming that all of the outlets from a tank are closed. This
capacity is to include relief from fire exposure.
The required venting capacity of maximum liquid movement into a
tank and the resulting vaporization should be equivalent to 0.159
cubic meters per hour for each 0.159 cubic meters (42 gallonbarrel) per hour of maximum filling rate. A liquid feed at or near the
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boiling point at tank pressure may require an outbreathing capacity
that is higher than the capacity indicated above. See API 2000
(Fourth Edition), paragraph 3.4.2.4 for further details.
o
Outbreathing resulting from maximum inflow of liquid into the tank
and maximum vaporization caused by such inflow.
The vacuum relief devices shall be suitable to relieve the flow
capacity determined for but not limited by the largest single
contingency or any reasonable and probable combination of
contingencies. It is permissible to reduce the requirement for
vacuum relief capacity by the rate of vaporization that results from
minimum normal heat gain of the contents. A gas-repressuring line
with a suitable control and source of gas may be provided to avoid
drawing air into the tank. If a gas-repressuring system is used, it
shall be used in addition to the vacuum relief devices, and no
capacity credit shall be allowed.
The requirement for venting capacity for maximum liquid movement
out of a tank should be equivalent to 0.34 cubic meters per hour of
air for each 0.159 cubic meters (42 gallon barrel) per hour of
maximum emptying rate for liquids of any flash point.
o
Outbreathing resulting from fire exposure
•
Single wall refrigerated storage
The equations for emergency venting for fire exposure for
single wall refrigerated storage tanks are given in API 2000
(Fourth Edition), paragraph 3.4.3.1.
•
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The heat input from a fire initially causes the vapors in the
space between the walls of a double-wall refrigerated storage
tank to expand, and the heat input also causes the vapors in
the roof space of a double-wall tank with suspended-deck
insulation to expand; however, it may be several hours before
the increased heat input to the stored liquid causes a
significantly increased vaporization rate. The venting
requirements for handling the increased vaporization may be
small compared to the requirements for handling the initial
volumetric expansion of the vapors.
Because emergency venting for a double-wall refrigerated
storage tank is complex, no calculation method has been
provided by API 2000. A thorough analysis of the fire relief for
a double-wall refrigerated storage tank should be conducted.
4.10.4 Means of Venting
4.10.4.1
Open vents
Open vents may be used to provide venting capacity for tanks in which
oil with a flash point of 100 °C (38 °C) or above is stored, for heated
tanks where the oil storage temperature is below the flash point, for tanks
with a capacity of less than 59.5 barrels (9.4 m3 ) used for the storage of
any product, and for tanks with a capacity of less than 3000 barrels (477
m3 ) used for storage of crude oil. The open vent is nothing more than a
hooded opening which is manufactured in several forms to connect to a
standard tank nozzle, equipped with a screen to prevent entry of birds,
insects, etc., and so shaped that it cannot become blocked with built-up
dirt or freezing moisture. Open vents allow free passage of vapor and air
in and out of the tank with changes in liquid level and thermal conditions,
and the requisite number to supply tank venting requirements is
provided. Most open vents can be supplied with flame arrestors if
required.
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4.10.4.2 Conservation Vents
Breather or conservation vents are required on tanks contents which have
a flash point of less than 100 °C (38 °C), per API 2000 (Fourth Edition).
Conservation vents are recommended for the following circumstances:
•
Tanks where the stored liquid is heated to or near the flash point.
•
Tanks where the flash point of the stored liquid is 86 to 140 °F (30 to
60 °C). This decision should be based on calculation of the cost
savings of reduced vent losses.
•
Tanks where inert gas blanketing is used to prevent air or moisture
contamination of the stored product.
4.10.4.3
Normal Venting
Normal venting shall be accomplished by an open vent or conservation
vent. See API 620 for further details.
If a pilot-operated relief valve is used, it shall be designed so that the main
valve will open automatically and will protect the tank in the event of
failure of the pilot valve or any essential part. Relief valves equipped with
a weight and lever are not recommended.
A tank that may be damaged by internal vacuum shall be provided with at
least one vacuum relief device set and sized to open at a partial vacuum
that is sufficient to protect the tank from damage.
A discussion of the types and operating characteristics of venting devices
can be found in Appendix C of API 2000 (Fourth Edition).
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4.10.4.4
Emergency Venting
Emergency venting may be accomplished by the use of the following:
•
•
Larger or additional relief devices
A gauge hatch that permits the cover to lift under abnormal internal
pressure
•
A manhole cover that lifts when exposed to abnormal internal
pressure.
4.11
REFERENCES
1.
Huff, J.E., “Frontiers in Pressure-Relief System Design”, Chemical Engineering
Progress, September, 1988, pp. 44-51.
2.
Grolmes, M.A., and Leung, J.C., Chemical Engineering Progress, Volume 81, pp.
39-46, December, 1985.
3.
Grolmes, M.A., and Leung, J.C., Chemical Engineering Progress, Volume 81, pp.
47-52, December, 1985.
4.
Grolmes, M.A., Leung, J.C., and Fauske, H.K., Chemical Engineering Progress,
Volume 81, pp. 57-62, December, 1985.
5.
Atherton, J., “Hydrocracker Explosion and Fire - BP Oil Grangemouth Refinery, 22
March, 1987”, presented at the 1988 Mid-Year Refining Meeting, American
Petroleum Institute, Refining Department, Operating Practices Committee, May 10,
1988.
6.
Chap4-r1.doc
Walker, J.J., “Sizing Relief Areas for Distillation”, Chemical Engineering Progress,
Pages 38-40, September 1970.
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7.
Bradford, M. and Durrett, D.G., “Avoiding Common Mistakes in Sizing Distillation
Safety Valves”, Chemical Engineering, Pages 78-84, July 1984.
8.
R.A. Crozier Jr., “Sizing Relief Valves for Fire Emergencies”, Chemical Engineering,
October 28, 1985.
9.
P.M. Brown, D.W. France, “How to Protect Air-Cooled Heat Exchangers Against
Overpressure”, Hydrocarbon Processing, August, 1975.
10.
Ernesto Valdes and Kenneth J. Svoboda, “Estimating Relief Loads for Thermally
Blocked-in Liquids”, Chemical Engineering, September 2, 1985, pp. 77-82.
11.
Fischer, H.G., et. al., Emergency Relief System Design Using DIERS Technology,
AIChE, New York, 1992.
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FIGURE 4.1
ISOTHERMAL FLOW OF COMPRESSIBLE FLUIDS
THROUGH PIPES AT HIGH PRESSURE DROPS
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The first requirement in designing a disposal system is to define the loadings to be handled.
Where the disposal system serves only a single pressure relief valve or depressuring valve,
the design basis for the disposal system would normally correspond to the design loading of
the pressure relief valve or to the maximum flow rate of the depressuring valve.
Where the disposal system serves more than one pressure relief or depressuring valve,
loadings should be calculated for all contingencies that may affect any of the pressure relief or
depressuring valves. This means that each pressure relief valve tied into the system should
have a flow rate, a temperature, and fluid properties calculated for each pertinent
contingency.
For disposal systems that serve more than one pressure relief or depressuring valve, the plot
location of each source should be defined. Combinations of sources into one or more
headers for disposal should be studied or defined. Based on the specified or assumed
combination arrangement, the loadings for each section of header and the rest of the disposal
system can then be defined. Maximum load is not necessarily the largest number of
kilograms per hour, but may be defined by criteria in either Chapter 7 (header system),
Chapter 8 (knockout and seal drums) or Chapter 9 (flare). For header system design, the
combination of flows that imposes the greatest head loss in flowing through the system is
controlling. For example, a flow of 110,000 lb/h (50,000 kg/h) of a vapor with a molecular
weight of 19 at a temperature of 300 °F (149 °C) develops a greater head loss and is a
greater load than a flow at 165,000 lb/h (75,000 kg/h) of a vapor with a molecular weight of 44
at a temperature of 100 °F (38 °C).
After major flare loads are determined, load mitigation options can be considered if calculated
loads exceed any of the disposal system limits in an existing system or to reduce costs in new
systems. Also, in many systems, it is desirable to reduce potential relief loads, within reason,
to allow room for future process unit expansion and to mitigate community concerns which
can arise from large releases.
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FLARE LOAD CALCULATIONS
5.2.1
General Methodology
To define the system load, the simultaneous occurrence of two or more unrelated
contingencies need not be assumed. For example, it is generally not necessary to
assume blocked outlets on process systems under fire conditions. The consideration
of two unrelated contingencies is commonly referred to as “double jeopardy”.
However, under some arrangements of process equipment, a fire could possibly result
in a failure of local wiring or instrument air piping, leading to the closure of valves that
block off the process system. Each individual contingency should therefore be
reviewed for possible resultant effects.
Cases of failure of major utilities, such as power or cooling water failure, require
particular study. The potential for complete loss of electrical power, cooling water, or
steam to an entire plant should be assessed. Where utility sources are believed to be
unreliable or are not backed up by a spare equipment, the effect of complete failure
should be studied. This type of study, with reference to electrical power failure,
commonly results in a design based on the failure of one bus or motor control center,
although loss of an entire distribution center or of the incoming line is occasionally
used as a basis for design.
Planned or emergency depressuring loads may add to the load from the relief devices
into the disposal system during some upset contingencies.
If the load is determined to be excessive based on disposal system considerations,
once a preliminary combined load has been established, then load mitigation methods
can be utilized to reduce those calculated loads. This analysis may include process
modifications, addition of instrumentation, or analytical methods to refine load
calculations.
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Determination of Relief System Loads
5.2.2.1
Determination of Area Fire Loads
To define the combined relieving loads under fire exposure, the probable
maximum extent of a fire should be estimated. This may be done on the
basis of the actual layout of facilities, considering the location of sources of
combustibles, the provision of drainage, and the effects of natural barriers.
Facilities that handle only gaseous fluids may be assumed to generate more
localized fires than those that handle liquid combustibles.
API RP 521 suggests that an equipment fire exposure area of about 2500 to
2
2
5000 ft (230 to 465 m ) be used as a basis for design. These are not firm
limits, but generally define the potential for multiple equipment items to be
involved by a single fire. Normally, the client specifies the fire exposure area
as a design guide in a project. However, if there is no guideline from the
2
2
client, a fire area of about 2700 ft (250 m ) (circle diameter of about 59 feet
or 18 meters) can normally be assumed for the purpose of flare load
calculations. However, if flammable liquids can be contained within an
enclosed area, then that entire area may need to be considered as a single
fire zone. A diked tankage area, for example, should usually be treated as a
single fire area.
Where equipment which contains liquid relieves along with equipment which
is vapor full through the same relief device, then the fire load should be
based only on the equipment containing liquid.
Latent heat and composition of the relieving material for common system
loads relieving through an individual relief valve can be based on a single
composition to simplify the calculations. For a reactor loop this composition
would typically be derived from the separator liquid composition and for
towers from the reflux.
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Utility Failure
Utility failure is often the determining factor for relief header sizing. Two
types of failure may occur: general or partial.
General failure is the more complicated of the two. Evaluation of the effect
of overpressure, attributable to the loss of a particular utility service, includes
the chain of events that could occur and the reaction time that is involved.
The consequences of a major shutdown must be considered.
For a partial utility failure, care must be exercised when taking credit for
equipment unaffected by the failure. In general, credit may be taken for
equipment in continuous parallel service having unrelated energy sources.
No credit may be taken for standby equipment, such as spare pumps or
compressors, even though the energy source may be unrelated. Manual
cut-in of auxiliaries is operator and time dependent and must be carefully
analyzed before being used as insurance against overpressure.
A.
Loss of Electric Power
Loss of electric power can cause a wide range of failure resulting in
overpressure. Power failure may be local, intermediate or total.
Local power failure affects the operation of individual equipment, such
as air fans, reflux pumps and solenoid valves.
Intermediate (partial) power failure affects the operation of one
electrical distribution center, one motor control center or one bus.
Depending upon how a group of motors are connected to the power
source, multiple failure conditions may occur. For example, assume
that the reflux pumps of a fractionator are motor driven and connected
to the same bus line. If a bus line failure occurs, the accumulator can
flood, eventually resulting in a flooded condenser. For improved
operability, the preferred practice would be to place motors in
complementary service on separate bus lines. For this configuration,
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the impact of loss of power from one bus will vary, depending on which
pump is operating at the time.
Total power failure is assumed to be 100 percent loss of electric power.
The effect of a power failure on all the flows in a unit must be evaluated
to establish the relief flows.
Backup electric power sources may be considered in the analysis of
the effects of electric power failure. For example, critical electronic
instrumentation is generally connected to the Uninterruptible Power
Supply (UPS) to ensure safe and orderly shutdown upon loss of normal
electric power supply.
B.
Loss of Cooling Water
Complete or partial loss of cooling water supply may be caused by
power failure or equipment breakdown. The layout of the cooling water
piping must be considered, as it may not be possible to have a
complete loss of cooling water to all units. The most common basis for
piping blockage involves loss of one lateral rather than the entire
supply header. Types of drivers on cooling water pumps must be
evaluated because standby pump drivers usually have an alternate
energy source.
C.
Loss of Steam
Steam or boiler feedwater failure may trigger a series of events that
could result in overpressure. The capability of steam systems to pick
up standby turbine loads should be reviewed in conjunction with the
overall installed boiler capacity and the normal standby capacity
immediately available. Overpressure may result from the loss of steam
to any turbine-driven reflux pump or cooling water pump.
In a series of fractionation columns where steam is the source of
reboiler heat, loss of steam to the reboiler of one column may cause
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overpressure in the following column. Loss of heat results in some of
the lower boiling point material mixing with the bottoms stream and
being transferred to the next column as feed. In this circumstance, the
overhead load of the second column may consist of its normal vapor
load plus the lower boiling point material from the first column. If the
second column does not have sufficient condensing capacity for the
additional vapor load, excessive pressure could occur.
D.
Loss of Fuel
E.
Loss of fuel may cause loss of steam, electric power and/or cooling
water, thus indirectly causing overpressure. The same comments as
for loss of steam to a series of fractionators apply to loss of fuel in
similar service.
Loss of Instrument Air
Instrument air failure may be local or general. Depending on the
service, local failure can cause overpressure for various reasons.
General failure causes plant shutdown with each control valve reverting
to its fail-safe position in conformance with safety and overpressure
limitation considerations. The fail-safe characteristics of each control
valve are established as an integral part of overall plant design. Failure
of the power supply to electronic or electrical instruments may also be
considered plant-wide unless proper standby power supplies are
provided.
5.2.2.3
Other Contingencies
While fire and utility loss typically result in the most significant relief loads,
maximum relief loads from individual relief valve services must also be
considered. These individual loads may control the sizing of a portion of the
disposal system piping. Less frequently, a large individual load may even
control the overall design of the disposal system. Large individual relief
service loads may result, for example, from loss of a reflux pump on a large
tower or blocked discharge on a major vapor stream.
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Loads from Depressuring Systems
Depressuring valves are normally assumed to be either 100 percent open or
100 percent closed. The maximum load from an open depressuring valve
normally corresponds to the flow capacity of the valve at the maximum
pressure of the protected equipment.
Where the same equipment is provided with both pressure relief valves and
depressuring valves, only the larger of the pressure relief valve load or the
depressuring valve capacity needs to be considered as the disposal system
load.
Should the capacity of a vapor depressuring valve exceed the normal vapor
flow rate within the protected equipment or if the depressuring rate is additive
to normal flows within the equipment, considerable liquid entrainment may
occur. Therefore, disposal systems for depressuring valves should generally
provide for liquid carryover.
Depressuring loads may occur simultaneously with relief loads due to a
common contingency. In some cases depressuring may be triggered
automatically, but most frequently will be implemented based on operating
procedures. The net impact of depressuring on the total flare load in those
situations must be considered.
5.3
FLARE LOAD MINIMIZATION
Flare load calculations must be based on sound engineering judgment, with full application of
local rules and regulations. Recommended practices and guidelines generally accepted by
industry should also be considered in arriving at an acceptably safe design.
Methods for calculating reductions in flare load from simple summation of all individual loads
can be broadly categorized into system modifications, methods based on qualitative
arguments and those which rely on risk-based statistical criteria or other quantitative
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calculations. Section 5.3.2 addresses system modifications; qualitative approaches are
considered in Section 5.3.3 and Section 5.3.4 is devoted to quantitative methods.
5.3.1
Background
Fail-safe devices, automatic start-up equipment, and other conventional control
instrumentation should not replace pressure relieving devices as protection for
individual process equipment. However, in the design of some components of a
relieving system, such as the blowdown header, flare, and flare tip, favorable
instrument response of some percentage of instrument systems can be assumed.
The percentage of favorable instrument responses is generally calculated based on
the amount of redundancy, maintenance schedules and other factors that affect
instrument reliability.
It is important to recognize that the concept of risk assessment ultimately must be
addressed in establishing a safe design. Use of a systematic method for addressing
risk provides consistency in the selection of a design basis. Without such a method,
differences in approach can develop between facilities and even between different
engineers working on the same design project.
Consistency with respect to risk assessment is required to assure all interested groups
(owners, employees, and local residents, for example) that risks are being properly
managed. Governments are also frequently involved in the effort to reduce the risk of
accidents. This is evidenced in the United States, for example, by the US
Occupational Health and Safety Administration regulation, "Process Safety
Management of Highly Hazardous Chemicals". As a result, having a consistent
documented technique for the selection and design of safety systems is not only good
management practice, but is evolving into a regulatory requirement.
Use of a consistent, systematic approach does not imply that quantitative risk
assessment is required. In fact, there are no universally recognized standards on
which to base a quantitative assessment. Under certain circumstances, quantitative
analysis is required or useful in determining the impact of an incident, for example,
calculation of maximum ground level concentration of a toxic compound resulting from
a release. More frequently, however, decisions are based on judgment, experience
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and qualitative (order of magnitude) estimates of risk, considering the likelihood that
an event will occur, the potential consequences of the event (both to health and
monetary), and the cost of any proposed method to mitigate the event (reduce its
frequency or limit the consequences).
In the following discussion, some of the approaches that are currently used by
different companies to assess or mitigate the magnitude of flare loads under major
failure scenarios are presented. It is assumed that in each case, the triggering event
will occur and that the maximum load based on worst case assumptions has been
calculated for each individual relief device. One issue addressed in Sections 5.3.3
through 5.3.4.4 is the extent to which the total release from multiple relief devices and
vents resulting from a single event may be reduced below that calculated by simply
summing the calculated loads from each device. In Section 5.3.4.5, consideration is
given to addressing the case where it is not practical to protect a system for a
particular upset condition with an individual relief device. This may be necessary, for
example, due to the magnitude of the load that would be generated or to the severe
impact of some upsets.
5.3.2
System Design and Modifications
A first step in addressing load reduction is to consider design alternatives which can
limit relief loads. These options are most easily incorporated into a new design,
thereby minimizing their cost impact.
Typically, the largest flare loads from multiple relief valves result from fire, power
failure or cooling water loss. System design modifications can be used to reduce
loads from each of these contingencies.
As discussed in Chapter 4, fire proofing insulation can be used to reduce relief loads
from equipment during a fire. Selective application of insulation to key equipment can
be used to reduce the overall fire load to the flare.
Particularly during the initial design stage, the electrical system can be configured to
help limit the impact of a single jeopardy power failure on the flare system. General
power failure, if a valid contingency, frequently is not the largest relief load
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contingency due to loss of key pumps and compressors which can reduce system
heat and mass throughput. However, local power loss can cause larger relief loads.
By carefully configuring the electrical distribution system, the impact of any single
power system outage can often be limited. For example, by using multiple
transformers and switchgear to supply 480 V power to an area instead of using a
single transformer system, the loss of a single transformer or bus will only impact a
fraction of the motors. By appropriate selection of driver power sources, relief loads
can be substantially reduced.
Loss of cooling water typically results from power loss shutting down motor driven
circulation pumps. By providing varied driver power source selection, either by using
different electrical buses for individual motor drivers or steam turbine drivers, any
reduction in cooling water circulation loss can often be moderated.
Other design options can also be applied on case specific bases to reduce or
eliminate significant flare loads. For example, design of a column for higher pressure
may eliminate reboiler heat transfer at relief conditions. Increasing the size of a
column accumulator may provide adequate surge volume to avoid flooding the
overhead condenser during an upset. Such design options, while not normally
economic, can be used selectively under the right circumstances to provide an
optimum economic solution, considering the economic impact on the flare system.
This can only be successful if design of the relief system is considered as an integral
part of the plant design.
5.3.3
Generalized Approaches
5.3.3.1
Percentage Reduction
The simplest approach in calculating flare loads is to assume that the actual
maximum load is a fixed percentage of the sum of the individual relief loads.
Load reductions would be based on experience. However, there are no
generally recognized bases supporting load reduction factors. In the
absence of this data, such an approach carries the drawback of not
addressing specific differences between designs which may mitigate or
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contribute to the load. As a result, this type of approach, although still in use,
is not widely accepted and can be difficult to defend.
5.3.3.2
Time Frame Analysis
Peak relief loads are not likely to occur simultaneously for all relief valves in a
system. By estimating the duration and the timing of the peak relief load for
each relief source, for a given upset condition, it is sometimes possible to
provide a conservative basis for relief load reduction.
This approach is most effective in addressing multiple relief modes which
result from the same incident, but which can occur during clearly distinct time
frames. Maximum relief loads resulting from reactor loop depressuring,
combined with other relief loads, can sometimes be addressed with this
approach. For example, during a unit fire which requires reactor loop
depressuring, operator response to initiate depressuring is expected to occur
within ten minutes. However, large vessels containing sub-cooled liquid,
which will result in the largest fire relief loads, may take considerably longer
to begin boil-off, particularly if they are insulated. As a result, it may be
possible to establish, with a high degree of confidence, that the peak loads
and possibly the majority of the loads from the two "sources" occur
independent of each other and should not be additive. This conclusion can
be reinforced by clear operating instructions which direct operations to begin
depressuring at the first indication of a unit fire.
A similar approach can sometimes be used when power failure causes tower
relief loads to occur and can eventually result in reactor loop depressuring. If
reactor loop depressuring is not triggered automatically by the event and the
system is not immediately threatened (depressuring is precautionary only),
then operating instructions which require that tower heat sources be reduced
prior to depressuring can be used to avoid significant overlap between the
depressuring and tower relief loads.
A similar argument is sometimes made to reduce loads occurring within a
tighter time frame, such as tower loads resulting from power failure. This is a
much more difficult argument to support since critical tower loads can be
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reached very quickly and by definition must occur before operator response
results in load mitigation. In the absence of detailed analysis, prediction of
the timing of these loads can be very difficult, since tower dynamics can
cause the loads to cycle. Even though very large relief loads may deplete
tower liquid inventory in a short time period, it is difficult in most cases to
define the exact timing of the load. However, based on a review of the
"calculated" pattern of several towers relieving due to the same incident, it
may be possible to establish some reduction in the peak relief load with
confidence. This approach is clearly far less valuable in this type of situation
than when the timing of different relief sources can be more clearly
differentiated.
Dynamic simulation can be used to determine the time dependent relief rate
and is discussed further in Section 5.3.6.
5.3.3.3
Response of Control Instruments
In evaluating the load from an individual relief device, the assumption is
made that no favorable response is provided by any control instrument to
reduce the relief load. If the normal response could act to reduce relief load,
then the control valve is assumed to remain in position.
In calculating the combined loads from multiple relief devices, it is possible to
consider relaxing this criterion to allow "normal" response of these control
valves. Favorable response could be incorporated where the control
direction is unambiguous. This would include for example single loop
temperature control on a tower reboiler or cascade loops where both
controllers would tend to act to decrease load. In the case of cascade loops,
the control mode with the least favorable response would be considered.
The drawback to this approach is the concern about response time of the
controls, which is the primary reason why favorable response is not
considered for individual valve sizing. Typically, response of controls is of
the greatest importance in relief analysis for distillation columns. Control
system response is tuned to relatively slow changes in key variables, in order
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to effectively control during normal operation. Otherwise sudden control
response could upset the tower during normal operation. Sudden reductions
in tower heat input as a result of normal control response, for example, are
unlikely in the short time frame that may be required to initiate tower relief.
This type of argument is more likely used as a qualitative basis to support
some percentage reduction of combined relief loads or to further support
decoupling of depressuring and tower relief loads, for example.
5.3.4
Risk Based Assessment
5.3.4.1
Risk Concept
Risk based assessment differs from the approaches described in the
previous section not only in associating a frequency (even if only semiquantitative) with mitigation techniques, but also in requiring adherence to
administrative controls. Just as relief system reliability depends on specific
maintenance and testing procedures and frequencies, risk-based mitigation
methods require clear documentation of the bases, adherence of the
operations and maintenance groups to these procedures, and strict
management controls to monitor conformance with the required procedures.
Risk for a process facility can be defined as a measure of potential loss both
in terms of the incident likelihood and the magnitude of the loss. This
couples an undesirable outcome (consequence) such as safety,
environmental or financial impacts, with the likelihood of that outcome. The
total assessment of these factors is usually considered in a HAZOP or a
similar risk assessment process. In this section, risk based mitigation
methods will be assessed based on their use to reduce the likelihood of an
adverse consequence occurring.
Specifically the methods discussed are intended to reduce the likelihood of
maximum flare loads occurring. One of the difficulties in addressing risk is to
define what is considered a level of acceptable risk.
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Although there is no universally accepted standard of risk (Ref. 1), there is a
general understanding that accidents with minor consequences can be
tolerated more frequently than accidents with catastrophic consequences. A
typical industry guideline is that an accident which threatens the lives of offsite individuals should occur with a predicted mean frequency of no more
4
than every 10 years. As an example of a conservative limit of design, the
US Department of Energy defines an accident as credible (worthy of design
consideration) if it could occur every 106 years. Because of the uncertainty in
our knowledge, calculations of expected frequency tend to involve some
speculation.
Another way to look at risk is to observe that current design standards for
relief valves pose an acceptable risk. Therefore, as long as the failure of the
relief valve is the dominant contributor to risk, the overall relief system design
is acceptable.
Relief valves, like any mechanical device with moving parts, are subject to
occasional failure. The rate of failure is very low when the relief valves are
properly specified, properly tested and maintained, and operating in a
favorable service. For example, safety valves in water or steam service in
nuclear power plants have the following average failure rates:
o
Failure to open
o
o
Premature opening
Failure to reclose
6
300 per 10 demands
6
3 per 10 demands
6
3000 per 10 demands
[Ref. 2]
Under favorable conditions, the average unreliability of a relief valve
(probability of not opening when needed) is (300/106), or 0.0003 per
demand.
Under the indicated favorable conditions, the average failure rate from one
plant to another varied by an order of magnitude. The reason for such
variation may be differences in the quality of design, operating conditions,
plant age, maintenance practices, and definitions of failure.
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Under unfavorable conditions, relief valves may fail much more frequently
due to poor specification (under-sizing or over-sizing), improper inlet or
discharge line sizing, improper testing or calibration, corrosion,
polymerization, or presence of debris in the process fluid.
For the purpose of discussion only, risk frequencies will be discussed in this
section based on events requiring ten relief valves to lift, unless mitigation
steps are employed.
5.3.4.2
Pump Driver Selection Philosophy
One question that frequently arises in design is how to influence relief
system design by selection of pump drivers. For purpose of discussion,
consider operation of reflux pumps on ten operating towers, since loss of
reflux typically results in major relief loads.
The primary goals of driver selection should be plant operability. If all electric
drivers are selected for both operating and spare pumps, then consideration
is frequently given to placing one pump in each service on a different
electrical bus than the other pump in the same service. By this means, a
local bus failure may only result in a brief unit upset and the unit can then be
operated off the other bus. A similar impact results if, for energy balance
purposes, one pump in each service could be driven by a steam turbine.
The impact of these two alternate systems on relief system evaluations can
also be assessed.
Assuming the loss of either bus represents a major flare load case, then the
most advantageous basis for the all electric drive case is to distribute
operating pumps evenly between the two buses, i.e., five of ten operating
pumps on each bus. In most cases, it would be necessary to consider that
as many as six or seven pumps out of ten could be operated from one bus to
provide operational flexibility. As a result, the flare load reduction value of
this option is relatively small, since much of the flare load could still occur if
the bus with most of the operating pumps were to be lost. Administrative
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controls would be required to assure that this bus load balancing assumption
is continuously maintained during the life of the plant.
In the case of the one motor drive and one turbine drive configuration, a
similar conclusion is reached except that the turbine drive pumps may
continue operating during a total power failure. However, since unit
operation is normally lost anyway during total power failure, this may not be a
significant advantage. Use of turbine drivers is also frequently subject to
change based on steam balance consideration and as a result any
administrative controls considered to provide a reduction in flare loads must
provide sufficient flexibility to allow for switching based on operational
considerations.
In summary, the selection of pump drivers and power supply system
configuration is normally based on plant operability considerations. Some
advantage can be taken for selected configurations to reduce major flare
loads. However, the reductions are typically modest and may not warrant
the reduction in operational flexibility resulting from the administrative
controls required.
5.3.4.3
Auto Start Spares
An assessment of the impact of auto start spares on flare load reduction can
be examined as an extension of the all electric case examined in Section 3.2.
Based on our definition of acceptable risk, and assuming an event impacts
10 relief valves, the probability of overloading the flare system should be less
10
than (1-(.9997) ) = 0.003, per loss-of-power incident. This value is based on
the premise that it is acceptable to assume that none of the ten relief valves
fails to lift, if none of the instrument (auto start) systems is present. If a
standby pump is de-energized, the operating pump will continue operating
and the associated relief valve will remain closed. But, if the operating pump
is de-energized, the associated relief valve will open unless the standby
pump is started. The number of relief valves that relieve into the flare
system equals the number of standby reflux pumps that fail to start after loss
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of power. A standby pump may fail for two reasons: 1) The standby pump is
locked-out for maintenance, or 2) There is a surprise failure of the pump,
motor, or control system.
The probability that a pump is locked-out is often termed the unavailability
and is equal to the average maintenance time required to restore the
component to service divided by the average interval between failures. The
control items associated with the pump have a negligible effect on
unavailability. The unavailability of a motor-driven pump can vary greatly
from one plant to another, depending on operating conditions, maintenance
practices, and logistical problems. The following are typical industry values
[Ref. 3, 4, 5], assuming that all preventative maintenance/inspection is
performed during plant shutdowns:
RESTORE
TIME, HR.
INTERVAL
HR.
UNAVAIL.
Refinery reflux pumps
27
33000
0.0008
GCC regenerator reflux
pumps
12
33000
0.0004
ITEM
If 24 hours of planned maintenance/inspection is performed every 6 months
(4380 hr), the unavailability increases by 0.0055, to a total of about 0.006 for a
typical plant. To allow for factors such as initial plant shake-down problems,
assume that 10% of the standby pumps are locked-out of service, in other
words, one pump.
Since the design reliability (or unreliability) is being compared to that of relief
valves, the relative comparison must be based on similar data. Nuclear power
data is used because it is more available and more carefully taken than
industrial data. To start a motor-driven standby pump, the following
components must function: overhead high-pressure or high temperature
switch, control signal wiring, control logic, motor starter, and motor-driven
pump. Human errors in opening or closing the suction and discharge valves
on the standby pump are detected by the operator on later shifts. The
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following failure rates [Ref. 2], which include typical human errors, can be used
for comparison to safety valve failure rates:
o
o
Pressure switch fails to operate
Temperature switch fails to operate
o
Wiring open circuit (per terminal)
6
0.1 per 10 demands
0.1
per 106 demands
5
per 106 hours
o
o
Control logic failure
Motor-driven pump fails to start
3
per 106 hours
5000 per 106 demands
Let us assume that failure of an operating component will be detected
promptly. Components which have demand-based failure rates are not
improved by testing. On the other hand, standby components with time-based
failure rates must be periodically tested to verify readiness. For demandbased components, the reliability per demand is one minus the failure rate.
For time-based components, reliability depends on the time since the last
performance test.
For a component that is regularly tested, the unreliability is the probability that
the component has failed since the last test. On average, the demand on a
tested component will occur halfway through the interval between tests. (See
Ref. 4 for further explanation of probability equations.) If each circuit has 4
terminals and is tested every month (720 hr), the probability of wiring failure
6
since the last test is ½(4) (720) (5) / (10 ) = 0.0072, and the reliability of a
single circuit wiring is (1 - 0.0072) = 0.9928. Likewise, the reliability of a single
6
logic controller is (1 - ½(720) (3) / 10 ) = 0.9989. The reliability of the standby
pump system is the product of the reliabilities of the process switch, wiring,
logic, and pump with motor and starter: (0.9999999) (0.9928) (0.9989) (0.995)
= 0.9867, and unreliability is 0.0133. The components that contribute most to
the pump unreliability (probability of surprise failure) are the motor and starter.
After loss of one side of the electrical system, anywhere from all to none of the
operating pumps will be de-energized. Ideally, the operators will try to maintain
about half of the operating reflux pumps on each side of the electrical system.
However, in the worst case, all 10 pumps will stop. One of the standby pumps
is assumed to be locked-out. The probability that all 9 of the remaining pumps
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9
will restart is (0.9867) = 0.8865. The probability that one or more pumps will
have a surprise failure is 0.1135, by difference. The probability that exactly
one pump will have a surprise failure is 9(0.0133) (0.9867)8) = 0.1075. There
are 36 ways for exactly 2 pumps to fail, so that probability is 36(0.0133)2
(0.9867)7 = 0.0058. Likewise, the probability of exactly 3 pumps failing is
84(0.0133)3 (0.9867)6 = 0.0002. The probability of 4 or more pumps failing is
negligible.
If the flare is designed for three relief valves discharging at the same time (one
standby pump locked-out and two more surprise failures), then the probability
of more pumps failing is shown to be 0.0002. This is significantly less than the
required maximum probability of overloading the flare system of 0.003.
Generic data comparing turbine-driven pumps to motor-driven pumps from
other facilities [Ref. 3] showed that the turbine-driven pumps were about 10
times more likely to fail to start. However, the value given for failure to start,
3
6
4.2 per 10 , is less than the 5000 per 10 [Ref. 2] used above for motor-driven
pumps. If turbine-driven pumps are chosen, their starting reliability and
availability in standby service should be evaluated. If the particular turbines
3
chosen are expected to have a failure to start rate greater than 5 per 10 , or if
the percentage expected to be unavailable as standby exceeds 10%, the flare
would have to accommodate the load from an additional relief valve.
Based on the problem described above, designing the flare system for the
simultaneous relief from three relief valves (out of the ten considered) after
loss of power poses an acceptable and justifiable risk.
In order to credit and maximize the potential reduction in flare load, it is
essential that the operating management realize that there should be a
balance between operating pumps energized from either of the two sides of
the electrical system, that the plant should be shut down if more than 10% of
the standby pumps are inoperable (or alternatively a more conservative basis
used, such as maximum 2 standby pumps out of service), and that regular
testing of the automatic starting controls is required to ensure acceptable risk
during operation.
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These tests should prove that seven out of nine available auto start pumps
always start on demand. If the test results do not meet this requirement, then
test frequency needs to be increased and/or more conservative flare load
criteria established so that the tests provide absolute confirmation of the
design basis.
It should also be noted that changing the test frequency from one month to six
months results in the need to consider the outage of an additional pump (four
out of ten pumps all together) in order to maintain a comparable level of overall
system reliability. This illustrates the need for an up-front commitment from all
groups involved to establish system design criteria which are practical and
which can be controlled.
5.3.4.4 Instrumentated Shutdown System
A reliable control system can be designed to be equal to or more reliable
than a relief valve in preventing overpressure for the same incidents. The
system could consist of pressure sensors, control logic and signal
transmission, and an actuated control valve which responds to an impending
over-pressure. The system must contain enough redundancy to ensure
acceptable reliability and should also result in acceptable risk to employees,
neighbors, and the environment. It is possible to meet any proposed risk
criteria using industrial standard controls, provided the control system is
regularly tested and maintained. Under unfavorable conditions, a control
system has the advantage that components can be located to avoid
corrosion, polymerization, or debris. In some processes (e.g., high-pressure
ethylene reactors) relief devices are not capable of responding fast enough
to prevent an explosive runaway reaction, so high-speed control systems
have become the accepted practice.
In crediting operation of an automatic shutdown system to reduce flare load,
reliability only addresses the shutdown action. The impact of the shutdown
action on flare loads must still be assessed separately.
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Double Loop
One control option is to use dual switches (pressure, temperature, or
other) to send signals via control logic and wiring to two solenoid
valves, either of which will cause the valve to move to the safe position.
If an operating control valve is also used to perform safety duty, and
the valve regularly opens/closes as part of normal operation, then there
is continuously assurance that the valve is ready to open/close on
demand. If each circuit has 4 terminals and is tested every month (720
hr), the probability of wiring failure before the next test is ½(4) (720) (5)
6
/ (10 ) = 0.0072, and the reliability of each single circuit wiring is
0.9928. Likewise, the reliability of a single logic controller is (1 - ½(720)
(3) / 106) = 0.9989, and the reliability of a single solenoid is one minus
the failure on demand (1-(0.999)) [Ref. 2] = 0.999. The reliability of a
single loop is the product of the switch, wiring, logic, and solenoid
reliabilities: (0.9999999) (0.9928) (0.9989) (0.999) = 0.9907, and the
unreliability is 0.0093. For two independent loops, the system
2
unreliability is the product of the individual loops (0.0093) = 0.0001.
So, this system reliability is 0.9999 per demand, which is greater than
the 0.9997 reliability of the relief valve. See Ref. 4 for explanation of
the probability relationships.
From this information, it can be shown that proposed redundant control
loops will have a greater reliability than a relief valve, or in other words,
a lower probability of over-pressure. Redundancy of control loop
components is needed to ensure reasonable reliability and ensure that
operation will not be interrupted when the standby portions of the
control loop are periodically tested. Controls should be designed to fail
safe in case of loss of 24V DC control power or loss of instrument air.
Since API recommended practice is to not replace individual relief
valve services with instrumentation, in general each relief valve will still
be designed to accommodate the maximum load, assuming the
instrumentation does not act to mitigate that load. Therefore, in
determining the relief load for a specific event, it will be assumed that at
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least one of the shutdown instrumentation systems fails to operate.
This is conservatively assumed to be the system which results in the
largest relief load. Based on the assumption of one failure out of ten
systems, the assumed test period could be relaxed to once every six
months without further impact, for the assumed configuration.
B)
Single Loop
In most units, there are existing single loop shutdown systems. As a
result, it is useful to address the credit that can be taken for such
systems in reducing calculated flare loads.
A comparison of the values for unreliability presented in Sections
5.3.4.3 and 5.3.4.4 shows that a single loop instrumented shutdown
system has an expected availability similar to that for an auto start
pump system. However, unlike an auto start pump system, it is
frequently not possible to test a single loop shutdown system. Further,
most existing simple shutdown systems are not subject to the strict
administrative controls and testing requirements required for
incorporation into the safety relief system design.
While the specifics of the systems must be reviewed, including the
frequency of testing, about four out of ten systems would be assumed
to fail in assessing flare load credit. No credit should be taken without
implementing the required administrative controls and test plan.
C)
Triple Loop
While a dual loop shutdown system provides a highly reliable approach
to reducing relief loads, some risk remains of inadvertent shutdown due
to a faulty signal. To minimize this risk, while retaining high shutdown
reliability on demand, a triple loop system utilizing a two out of three
voting system is sometimes used (see Figure 5.1). Application of
voting logic is most reliably accomplished using a Triple-Modular
Redundant (TMR) control architecture, instead of PLCs. TMR employs
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three isolated parallel control systems and extensive diagnostics
integrated into one system. Two-out-of three voting logic is used to
provide high integrity operation with no single point of failure. A typical
application of this type of system is for burner safety management in a
fired heater. Shutdown/safety instrumentation can be integrated into
the burner management system configuration. A typical TMR system
with multiple shutdown systems is almost equal in cost to the
equivalent dual loop systems configured with DCS. The need for this
additional instrumentation should be assessed on a case by case
basis.
5.3.4.5
High Integrity Protective Instrumentation Systems
Some relieving scenarios and/or safety systems require the installation of
high integrity protective instruments systems to prevent over pressure and/or
over temperature. This occurs more frequently in the chemicals industries
than in refining, particularly in reaction systems which can run away more
rapidly than a normal pressure relief system can protect against. In these
situations, combinations of instrumented systems are used to minimize the
potential for an event to occur.
If this approach is used, the protective instrument system shall be at least as
reliable as a pressure relief device system, and shall be used only where the
application of pressure relief devices is impractical. Application of this
system implies that risk is significant and that the emphasis should be placed
on installing a high integrity system.
Due to the risk based decision to use an instrumented system to eliminate a
significant individual relief case, client’s management and operations
approval should be obtained for each system. Further, specific system
reliability should be estimated, documented, reviewed and approved. Where
practical, a standardized approach shall be used to minimize the effort
required to produce documentation and obtain approval for individual cases.
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As an illustration, a fired heater may supply so much heat to a tower feed or
bottoms that continued firing of the heater during an upset condition resulting
in loss of cooling can produce an enormous relief load. While this is
qualitatively different from designing to avoid runaway reactions, it can be
argued that "over-design" of a flare system to accommodate such a load
carries implied safety risks in addition to significant costs. This decision of
relative risk properly rests with the client.
Mitigation of the load from such a system should properly include added
reliability for the instrumented system due to the high implied risk level. To
obtain suitable reliability for such an application, redundant, independent
instrumented shutdown systems with at least double loops are
recommended. As an example, this could include a dual pressure
switch/solenoid shutdown of the fuel control valve to the heater and a
separate electrical bus voltage dual sensor switch trip linked to an
independent heater fuel shutdown valve. More frequently, due to the critical
nature of these trips and a concern for avoiding inadvertent system loss,
TMR shutdown systems are provided. Such a system can also provide
critical protection to the heater tubes in the event of loss of the heater charge
pump which could not be protected by a pressure relief device alone in this
circumstance.
5.3.5
Recommendations for Relief Load Mitigation Instrumentation
5.3.5.1
General
API recommended practice 521 allows for favorable instrument response of
some percentage of instrument systems in the design of the blowdown
header, flare and flare tip. The percentage of favorable instrument
responses is generally calculated based on the amount of redundancy,
maintenance schedules, and other factors that affect instrument reliability.
However, no specific recommendations are made for applying this concept
uniformly.
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In practice, no two companies appear to apply the same methodology in
crediting instrument response to reduce overall flare load. Approaches used
range from simple percentage reduction of flare load without regard to
specific reduction factors to statistical analysis of load probabilities.
There are many factors used in calculating loads from individual relief valves
that are assumed to be aligned to present a worst case condition. The
likelihood of these factors all aligning in this worst case condition for several
relief valves simultaneously is small. However, the impact of each of these
factors vary and it would be difficult to correctly analyze the individual and
combined impact of each of these factors. Further, it is intended that the
flare design limits not be exceeded to a high degree of certainty. For the
purpose of evaluation, this degree of certainty is assumed to be similar to
that intended for relief valve performance to protect equipment. On that
basis, many of the simple assumptions, such as selection of which pumps
are operating when one of two parallel electrical buses are lost, assume very
little significance, since the likelihood of these occurring are many times
greater than the limit of certainty required for the system.
Of key importance in crediting instrument response is clear documentation of
those instruments as part of the relief system and verification of the assumed
reliability of the instrument response. Any change in this instrumentation
must be carefully tracked and reviewed prior to implementation to assure any
impacts on safety are identified.
5.3.5.2
Normal Control Response
For calculation of individual relief valve loads, no credit is taken for normal
control response which tends to reduce the calculated relief loads. This
limitation does not apply in calculating flare loads resulting from general
failures which result in multiple relief valves lifting.
A key reason for not crediting normal instrument response favoring lower
relief loads for individual relief valves is that response time is difficult to
predict. Operating controls are set up to provide measured control to
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parameters that normally change relatively slowly compared to changes
which occur during upset conditions. This same concern still applies during
a general failure condition, but if carefully addressed for a specific situation it
should be possible to allow appropriate credit for control response.
Credit can be selectively applied for normal instrument control response
which tends to reduce flare load due to a general failure condition, based on
the following criteria:
A)
Credit should only be taken if the control action is direct and
unambiguous. If cascade control is used, both controllers must tend to
reduce flare load for any credit to be taken and the controller action
resulting in the greater relief load should be considered to be
controlling.
B)
A careful review of the controller response time relative to the impact
on the relief condition must be made to assure that the response credit
is appropriate to that time frame. Client must review the response time
considered and agree that there are no likely changes that will be
made or other existing factors that will cause that response time to be
exceeded.
C)
Inclusion of the normal instrument response as part of the relief system
design must be carefully documented. Any proposed changes which
impact that response must be carefully analyzed and any impacts on
the relief system design addressed before those changes are
implemented.
D)
Added caution should be used during review of controller response for
any control valve which is not normally modulating (i.e. is either
completely closed or wide open), since these valves will have a greater
tendency to be stuck in position and this condition may not be detected
during normal operation.
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Based on typical system design, these criteria will preclude crediting most of
the control instrumentation, since heat sources to most tower reboilers are
supplied on flow control, cascaded to a master controller. However, for
selective situations, following careful review, some normal instrument
response can be credited.
5.3.5.3
Non-Normal Automatic Instrumentation
While advantage can be taken of system reliability factors to also reduce
flare load as indicated by the discussions in the previous sections, inclusion
of specific shutdown systems is clearly the most effective approach to
minimizing flare loads resulting from relief of multiple tower systems.
Instrumented shutdowns provide a verifiable and reliable approach to
significantly reduce flare loads for this type of system. Such systems have
become a commonly used and accepted method for flare load reduction.
Auto start and auto lockout instrumentation is designed to function only
during upset conditions. The reliability of these systems depends strongly on
the degree of redundancy provided, the reliability of the individual system
components, and the frequency of system testing. Unlike normal operating
instruments, in most cases it is not obvious if some components have failed
in these systems, unless specific testing is performed. However, with
modern instrumentation, it is possible to install systems that are largely selfchecking.
For purpose of discussion, it will be assumed that all types of instrumented
systems will be tested at the same intervals. As a reference point this
interval is assumed to be 6 months, which is a practically long duration that
also provides room to be increased or decreased based on actual test
experience.
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Single Loop Systems
For the typical pump auto start system discussed in Section 5.3.4.3,
increasing the test period from one to 6 months results in the following
component reliability values:
o
Process switch
0.9999999
o
o
o
Wiring
Logic
Pump w/motor & starter
0.9568
0.9934
0.995
On that basis the overall reliability is: (0.9999999) (.9568) (.9934)
(.995) = .9457 . To provide a basis for crediting auto start pumps,
assuming the same number of instrumented systems and PSVs, the
number of instrument malfunctions is set so that loss of another
instrument is less likely to occur than malfunction (failure to lift) of one
of the PSVs. For example, the chance of one of 5 PSVs not lifting on
demand is (.0003) (5) = .0015 . If one out of 5 auto starts is assumed
to not work, the chance of a greater number failing is .017 . However,
if two out of 5 auto starts are assumed to not work, the chance of a
greater number failing is less only .0008, which is less than .0015 .
Therefore, two out of the 5 auto starts are assumed not to work on
demand during an upset. In addition, as discussed in Section 5.3.4.3,
one pump is assumed to be locked out and therefore not even
available. This leads to the assessment that for 6 auto start systems,
only 3 should be considered to work on demand during an upset.
Table 5.1 summarizes the calculated extent that credit should be
allowed for varying numbers of auto start pumps. Normal conservative
practice is that pumps are considered in order of descending impact on
total relief load (largest relief load impact considered first).
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TABLE 5.1
PUMP AUTOSTART LOAD REDUCTION CREDITS
PUMP #
CREDIT
Yes/No
1
No
2
No
3
No
4
Yes
5
Yes
6
Yes
7
No
8
Yes
9
Yes
10
Yes
While it is possible to take credit for these systems there are several
drawbacks, including the following:
o
As indicated above, no credit is considered for the 3 most
significant load reduction options. Obtainable credit may be
substantially reduced as a result.
o
The impact of auto start pumps may be specific to a limited
number of cases. For example loss of all 480 V power in an
area may cut power to both the normal operating and spare
pumps, so that an auto start spare has no impact on the load
for that case.
o
Even where an auto start pump reduces the relief load, it may
not eliminate relief. For example, auto start of a reflux pump on
partial loss of power which shuts down the normal operating
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pump may eliminate condenser flooding and total loss of
condenser capacity, but relief may still occur due to reduced air
condenser duty because of fan loss.
o
Due to the above considerations, the calculated load impact per
instrument system may be relatively small. Inclusion of many
instrument systems into the safety system design will add to the
administrative and maintenance burdens.
As discussed in Section 5.3.4.4, the reliability of a properly designed
single loop shutdown system is similar to that for a pump auto start
system. Therefore, the credit table provided for pump auto starts
applies also to single loop shutdown instrumentation. Since shutdown
devices are typically targeted to remove heat input to a tower system,
successful operation of one of these devices usually has a more
significant impact on flare loads than successful operation of an auto
start pump. As a result, some of the drawbacks noted for the auto start
instrument systems do not apply as strongly to shutdown systems.
However, the need to assume that the 3 shutdown systems with the
biggest flare load reduction potential do not function on demand,
remains a serious drawback to this approach.
B)
Multi-Loop Systems
As noted in Section 5.3.4.4, a well designed double loop shutdown
system can have sufficient reliability to assume that only one out of ten
shutdown systems does not function on demand, based on a six month
testing interval. However, existing shutdown systems are already
configured in many different ways. While in the future it will be
desirable to standardize the basic configuration used, it is appropriate
to base reliability on a typical installed system.
As an example, consider a typical instrument configuration used to shut
down a heater on loss of flow from the reboiler circulating pumps. This
system has dual sensors, a single logic block and dual shutoff valves
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(double block and bleed), each with dedicated solenoid valves for both
the fuel oil and fuel gas streams. System reliability is the product of the
input, logic and output reliabilities. Input reliability is high due to the
redundant switches. Output reliability is also high, even though two
streams must be shut down. As a result, the limiting factor is the
reliability of the single logic block. Based on a six month test interval,
the following reliability factors apply:
o
o
o
o
Input
Logic block
Fuel gas output
Fuel oil output
0.9981
0.9934
0.9981
0.9981
On that basis, the overall reliability is: (0.9981) (0.9934) (0.9981)
(0.9981) = 0.9877 . Using the same rationale for crediting instrument
response as for the auto start instrumentation, Table 5.2 indicates
credits which can be taken:
TABLE 5.2
DUAL LOOP SHUTDOWN SYSTEM
LOAD REDUCTION CREDITS
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S/D INSTRUMENT #
CREDIT
Yes/No
1
No
2
Yes
3
Yes
4
Yes
5
Yes
6
No
7
Yes
8
Yes
9
Yes
10
Yes
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This allows credit for 4 of the 5 instrument systems having the greatest
impact on relief loads. In most situations this will provide most of the
load reduction which can be obtained through instrumentation. This
method of applying instrumentation to reduce major relief loads
eliminates the concerns indicated for the single loop systems and is
therefore the preferred approach for instrumented flare load reduction.
Caution should be used in applying this approach with existing systems
to assure the needed redundancy is provided for high reliability. For
example, a typical heater shutdown system recommended by UOP for
a similar system may only include a single shutdown valve and
solenoid for each stream. While this may seem at first glance to be a
comparable shutdown system, in fact the reliability of that design would
be lower than that for a single loop auto start system. Sensing
instrumentation can be of many varieties, including for example
instruments that can detect power loss, pressure increase, temperature
increase, or flow loss, as long as dual, independent sensors and
signals are provided. Upgrading of the shutdown system to incorporate
a TMR configuration with two-out-of three voting logic would allow
credit for 9 of 10 instrument systems to function on demand.
This type of approach makes sense for a heater system which may
lose air fans during a power outage, for example, since heater
protection is of significant concern.
For most applications, for the reasons stated above, a reliable multiloop trip system is recommended for instrumented flare load reduction.
This is by far the most widely used methodology by refiners to
accomplish load reduction via instrumentation. If for a particular
application, it is also desired to credit some single loop instrumentation,
similar to that already described, then it is recommended that credits
be identified first for the reliable multi-loop systems and then the single
loop criteria should be applied using the modified flare loads as a
starting basis.
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High Integrity Protective Instrument Systems
Section 5.3.4.5 describes the criteria to be applied in designing a
system which actually eliminates an individual relief load case from
consideration. Due to the critical nature of these applications, TMR
shutdown systems are provided with two-out-of three voting logic. It is
desirable to standardize both the overall system design and the
individual components used for such systems to the maximum extent
practical. In refinery applications, such systems are most commonly
utilized around fired heaters. These systems are typically integrated
with the overall heater control instrumentation and logic. A system
sketch of a high integrity shutdown system is provided as Figure 5.1.
5.3.6
Dynamic Simulation
Dynamic simulation can determine realistic flare header flowrates which may help
avoid an expensive flare or flare header replacement. Conventional methods for
calculating flare loads are conservative and, in some circumstances, yield high flare
header back pressures. Dynamic simulation can complement other flare load
reduction methods such as the use of automatic shutdown systems.
5.3.6.1
Individual Column Relief Loads
Distillation columns are the largest contributor to the total refinery flare load.
However, only key distillation columns should be analyzed using dynamic
simulation to determine their relief load as rigorous dynamic simulation of
column relief is currently a time-consuming activity. Generally, dynamic
simulation should be used only after conventional design approaches have
identified the need for further flare load reduction. Dynamic simulation for an
individual column should be based on the column’s contribution to the total
flare relief load. Good candidates for dynamic simulation include crude and
vacuum columns, FCC fractionators, and hydroprocessing unit fractionators.
An example of the application of dynamic simulation to tower relief load
analysis is provided in Appendix B-7. The following benefits should be
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considered when assessing the potential for dynamic simulation for a
particular column:
o
Conventional methods may overpredict the relief load for columns with
wide boiling range feeds by a factor or two. Dynamic simulation
reduces the calculated relief loads for these columns as it accounts for
the limited inventory of light components and the sensible heat required
to heat the fluid to boil heavier components.
o
Dynamic simulation can provide a realistic calculation of the reboiler
pinch at elevated pressures using the time-dependent reboiler feed
composition.
o
Dynamic simulation can provide an accurate prediction of the time
required to pressurize a column before the relief valve lifts to assess
the potential for operator intervention.
o
Dynamic simulation provides insight into column behavior during relief
conditions. It can be used to make control logic decisions such as it is
better to maintain or stop reflux after a cooling water failure.
When dynamic simulation is used for process equipment and process safety
design, it is necessary to ensure that all the assumptions used in the model
are conservative, within the context of a given relief scenario. In this
manner, if dynamic simulation is used to calculate the relief flow from a
column, the highest calculated flow may be used to size the relief valve and
may be used in a total flare load analysis. If all the assumptions used in the
dynamic simulation are conservative, the actual relief flow will not exceed the
calculated flow if the contingency which was simulated occurs (except for the
instantaneous flow due to oversized relief valves). Despite this conservative
approach, equipment design conditions calculated by dynamic simulation
may be less severe than the conditions determined by conventional
calculation methods.
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Timing of Individual Relief Events
Total flare loads are typically calculated assuming all relief sources relieve at
their maximum rate simultaneously. Dynamic simulation can be used to
determine the time dependent relief rates for individual columns to determine
a time dependent total flare load. However, as discussed in Section 5.3.3.2,
careful consideration must be applied before superimposing the dynamic
relief flow profiles of several columns to reduce the total flare load. Many
unexpected external events can alter the timing of relief flows from individual
columns. These events may be caused by operator intervention, the failure
of automatic shutdown systems, cycling of relief valves, or other factors
which are not possible to foresee and to incorporate into a dynamic
simulation study.
Furthermore, a dynamic simulation that was used to determine the maximum
relief flow for an individual column will have incorporated assumptions that
provide a conservative upper limit for the column relief load, but may provide
an unreliable estimate for the timing of the relief flow. An example
supporting this is the duty associated with an air cooler condenser.
Assuming twenty-five percent duty in the condenser after a power loss is
conservative and leads to a fast column pressurization. However, the air
cooler may provide substantially greater cooling at elevated temperatures
found during relief conditions which will reduce the rate of pressurization. If
the condenser floods both the twenty-five percent case and the higher duty
case may provide similar relief rates, but at different times. Consequently, if
conservative assumptions are used to calculate an individual column relief
load, the actual timing of the relief may be different than the timing prediction
by the dynamic simulation.
Therefore, thoughtful consideration must be used when applying the timing
of relief events determined by dynamic simulation. A separate dynamic
simulation may be needed to be performed with a set of assumptions which
is conservative for the total relief load and not for the individual column relief
load. The number of relieving sources calculated by dynamic simulation
must be limited to those which can be studied in a manner which insures that
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the entire approach yields a conservative total flare load. Most likely this
number may be two or perhaps three dynamically calculated column relief
loads. All other relief loads must be assumed constant with time at their
maximum relief rate.
5.3.6.3
Flare Packing
Dynamic simulation can account for pressurizing (packing) the flare or
blowdown system. Packing reduces flare loads for events such as
hydroprocessor depressurization where it is known how the flare load
decreases with time.
5.3.6.4
Stage Depressurization
Dynamic simulation can be used to evaluate staged relief systems. In some
systems, such as high pressure compressor stations, the equipment is
depressured after an emergency shutdown in stages. This approach allows
for a smaller flare. Dynamic simulation can be used to test the system to
verify that it is properly designed.
5.3.7
Probability Analysis
A similar, but more time consuming approach to the risk analysis discussed in Section
5.3.1, is to complete a probability analysis addressing the likelihood of peak flare loads
occurring. Mitigation systems are analyzed to determine their reliability (or
unreliability) and flare load rate reduction potential if the system successfully activates.
Based on these inputs, a flare load probability distribution is generated and the
maximum credible relief load is selected as the relief load on the distribution
corresponding with an acceptable limiting occurrence frequency (for example once in
10,000 years).
One advantage of this approach is that if all the data is properly considered, all factors
which tend to reduce flare loads can be considered without limiting their impact
through simplifying assumptions. For example, it is not necessary to assume that the
"maximum" impact instrument fails to operate first, since there is a probability
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assigned to that occurrence. As a result, the conclusions of a thorough probability
analysis will be less conservative than the methodology discussed in Section 5.3.1.
There are also potential drawbacks to this approach. The maximum advantage of this
approach is obtained for reductions based on the operation of many instruments. This
both increases the extent of analysis required and incorporates more instrumentation
into the relief system maintenance requirements. Future modifications to the system
may require similar analysis, which in the extreme may require the original contractor
to reanalyze the system if the numerical basis and analysis are extensive.
While this approach is valid, careful consideration should be given before applying it to
a specific flare system. In many cases, an adequate solution can be reached through
application of process changes, high reliability instrumentation systems, or other
methodologies discussed earlier in Section 5.3. Clear application of a limited number
of mitigation steps is preferred, since this simplifies the future analysis, maintenance
and tracking of any added system components.
5.4
REFERENCES
1.
J.O. Philley, "Acceptable Risk -An Overview," Plant Operations Progress, October,
1992, pp. 218-223.
2.
S.A. Eide, S.V. Chmielewski, and T.D. Swantz, "Generic Component Failure Data
Base for Light Water and Liquid Sodium Reactors PRAs," EG&G Idaho, Inc., Idaho
Falls, 1990.
3.
"Guidelines for Process Equipment Reliability Data," American Institute of Chemical
Engineers, New York, 1989, pp. 37, 192, 195.
4.
M. Modarres, "What Every Engineer Should Know About Reliability and Risk
Analysis," Marcel Dekker, Inc., New York, 1993, pp. 24, 128-135, 200-206.
5.
R.P. Dawkins and J.A. Derdiger, "Component Failure and Repair Data: GasificationCombined-Cycle Power Generation Units," Electric Power Research Institute, Palo
Alto, 1982, pp. 2-98.
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Cassala, J.R., Dusupta, S., and Gandhij S., “Modeling of Tower Relief Analysis”,
Hydrocarbon Processing, October 1993, p. 71, and November 1993, p. 69.
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TRIPLE LOOP SHUTDOWN SYSTEM
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Pressure relief streams must be disposed of in a safe, economic and environmentally
acceptable manner. The basic concerns that should be considered in defining disposal
locations and methods for pressure relief and other emergency release streams are
described in this section. Due to the diversity of types of streams, potential release
quantities, local regulations and existing facilities, the discussion is generally limited to
criteria which should be considered in defining release locations rather than defining
specific practices.
Each project should evaluate the characteristics of the emergency releases that may occur
and available disposal alternatives and develop a plan for emergency release disposal.
This plan should identify the major potential releases and describe where each stream will
be routed for disposal. This plan should be reviewed with site management as a part of
project scope definition and should be made as early as possible in a project, preferably
during the conceptual phases, so the costs associated with disposal of emergency release
streams can be included in cost estimates and economic assessments.
6.2
DISPOSAL OPTIONS
There are several common alternatives for disposal of pressure relief streams. Each of
these has its own economic, safety and environmental advantages and disadvantages.
However, the discharge of flammable or toxic vapors to atmosphere should be minimized.
The criteria for evaluating the suitability of these alternatives are discussed in the following
sections.
The basic alternatives are:
6.2.1
Discharge to Atmosphere
This method routes the release stream directly to the atmosphere and relies upon
natural dispersion and jet mixing to reduce the concentration of released material to
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safe levels. This alternative is often the most economic method, and for many
streams is as safe or safer than other methods. However the safety and
environmental aspects of a release must be assessed. This is discussed in Section
6.5.
6.2.2
Discharge to Grade or Sewer
This method routes the release stream directly to the ground or to a sewer system.
The released material is recovered in a sewer treating system or must be cleaned
up after a release. This method can generally be used for liquids which are nonhazardous and are not released in large quantities. Criteria and considerations for
this type of release are discussed in Section 6.6.
6.2.3
Discharge to a Process Vessel
This method returns a released fluid back into a process system operating at lower
pressure than the system from which it was released. This method is attractive for
streams that are unsuitable for grade or atmospheric release and usually does not
require construction of collection or treating systems. The impact of doing this on
the receiving process must be evaluated. This is discussed in Section 6.7.
6.2.4
Discharge to a Closed Collection System
This method discharges released fluids to a closed collection system. This system
may treat or cool the release stream and either recover some or all of the material or
route it to a remote location where it can be safely disposed of. This type of system
can be a liquid blowdown system, an incinerator system, a flare system, a burn pit,
or a vapor recovery system. These systems are discussed in Section 6.8.
6.3
HAZARD AND RISK ASSESSMENT
Decisions on where to dispose of pressure relief release streams require assessment of the
hazards and risks associated with various disposal alternatives. Some of the
considerations in determining acceptable release criteria are:
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o
Environmental requirements or restrictions which might apply to a release.
o
The toxicity of the potential release streams.
o
The volatility and flammability of the stream.
o
The molecular weight of a vapor release stream.
o
The frequency and potential quantity of release.
o
The number of simultaneous failures which would have to occur for a release to be
considered hazardous.
Most releases from pressure relief systems will occur only during emergencies, and
therefore, the same criteria that might be applied to continuous or routine releases
generally do not apply. This is particularly true when environmental regulations are
considered. Most environmental restrictions either do not apply or make special provisions
for releases which are classified as emergency releases. However, emergency releases
generally must be included in total emissions reports to regulatory agencies.
In considering relief stream disposal, it also should be recognized that these releases are
most likely to occur during emergencies when a process may be partially or completely out
of control. The disposal methods selected must consider this possibility and not overly rely
upon normal process conditions, normal operation of equipment or control systems or the
availability of manual intervention.
6.4
ENVIRONMENTAL FACTORS
As mentioned above, most pressure relief system releases are considered to be emergency
releases and are not subject to the same environmental restrictions as continuous or
routine releases. However, the environmental aspects of emergency releases from
pressure relief systems must be considered. Among the factors which should be
considered are:
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Generally the total releases from a facility, including emergency releases are
regulated. Estimates of total yearly releases generally must be reported to
regulatory organization.
o
Releases of certain types of toxic or hazardous materials are of specific concern.
Some of these materials which are typically found in refineries, which fall into this
category, are listed in Table 6.1.
o
Public relations concerns on how a facility is perceived by its neighbors are
important. Releases that may be acceptable from a legal standpoint may not be
acceptable to the local community.
All of these concerns are beyond the scope of any one project or installation. Disposal
practices based upon these concerns must be defined by site management.
6.5
VAPOR RELEASE CRITERIA
Many vapor relief streams may be safely and economically disposed of directly to
atmosphere. API RP 521, paragraph 4.3 discusses the characteristics of vapor releases,
potential hazards which should be recognized and design considerations. Some of this
material is summarized below. However, an engineer determining the suitability of a stream
for atmospheric release should be familiar with the discussions in API RP 521 and local
regulations.
6.5.1
General
Two basic criteria apply when evaluating a pressure relief stream for suitability for
atmospheric release:
o
The release of the stream should not produce an immediate hazard to
operating personnel or people or facilities in the vicinity of the refinery.
o
The frequency and quantity of the release should be sufficiently low so that
the amount of pollution or risk of damage to property internal or external to
the facility is acceptable.
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Atmospheric Release Criteria
There is no universally accepted set of criteria for establishing when a given relief
stream is suitable for atmospheric release. There is very little hard data to provide a
basis for firm guidelines, and the determination of whether a stream may or may not
be released is often a subjective decision depending upon an individual’s training
and experience. Fortunately, the refining industry has not had a catastrophic loss
that could be traced to atmospheric release of a vapor relief stream.
Although experience has been good with atmospheric releases, minor incidents and
potentially hazardous situations have occurred. Therefore, the suitability of the
streams for atmospheric release should not be assumed, but should be reviewed for
each installation. Some characteristics which indicate that a stream is suitable for
atmospheric release, provided that no other hazards will be produced, are listed
below:
o
The release must be a vapor and be from a vapor space of a vessel that
affords sufficient disengaging area that release of liquid or liquid entrainment
by the vapor release is not a concern.
o
The release must be from a pressure device which will open only under
emergency conditions.
o
The release must not be a toxic substance. This is intended to mean a
substance which is immediately toxic such as hydrogen sulfide or similar
materials. Some releases of “toxic” substances may be acceptable if a
dispersion modeling calculation shows that the concentration will be reduced
to safe levels by jet mixing before any personnel can be exposed to the
release. Additionally, environmental regulations may limit the frequency or
quantity of the release of other materials which are not directly toxic, but
considered to have long term health effects.
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The released vapor should not condense or form a mist or droplets that may
settle back to grade. There is no firm criteria for determining this, but
releases of vapors which are heavier than air are often considered to be
candidates. The behavior is strongly affected by atmospheric conditions and
is very difficult to predict.
o
The total quantity of release should be small enough that formation of
hazardous vapor clouds will not occur.
In addition to the characteristics of the potential release, other factors relative to the
installation should be considered. Location of the release in a congested area or in
areas where public facilities are adjacent to the release may indicate that other
disposal methods should be considered. Ambient conditions which may exist also
must be assessed. Ambient temperatures or prevailing winds can also affect a
decision on the suitability of a stream for atmospheric release.
6.5.3
Safety Review
Releases, which may not meet the above criteria, should be reviewed with site
management before a decision is made to allow atmospheric release. Some of the
criteria that indicate that a higher level review is needed are:
o
The stream has the potential for being liquid or bearing significant amounts
of liquid. Streams from vessels which may fill up with liquid or which could
entrain significant liquid (such as might occur in a heavily loaded
fractionation tower) should be reviewed for suitability.
o
The stream has the potential to condense or settle back to grade before
being dispersed. High molecular weight vapors or any vapor heavier than air
that might be released in large quantities should be reviewed.
o
The release stream is hot, corrosive, toxic or otherwise hazardous.
Decisions on determining which streams are suitable for atmospheric release are
qualitative and must be based upon knowledge of the material which may be
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released, its potential hazard creating properties, experience with similar services,
specific site related considerations and current laws and regulations. The final
decision is usually one based upon hazard and risk assessment by site
management.
6.6
LIQUID RELEASE CRITERIA
Liquids should never be released to atmosphere, but in some cases liquid releases from
pressure relief devices may be directed to a safe containment area at ground level. The
following criteria should be applied in assessing whether liquid release streams may be
routed to locations other than a closed system.
6.6.1
Non-Hazardous Streams
Cool or warm water streams [less than 140 °F (60 °C)] may be directed to grade
irrespective of quantity. Hot water or streams that may flash should be routed to a
suitable flash drum.
Non-toxic hydrocarbon streams from pressure relief valves installed for liquid
thermal relief may be routed to grade or a sewer. These should be true thermal
relief where a sustained flow is not expected and other relief cases for these valves
(i.e. fire, blocked discharge, etc.) do not exist.
6.6.2
Non-Hazardous Hydrocarbons
Non-volatile, non-toxic liquid streams such as cool, heavy hydrocarbons should be
contained in a closed system wherever practical. If containment of a non-volatile
stream is not practical, it is often permissible to route the stream to an oily water
sewer or a safe open containment area near the point of discharge. Such a
discharge should not create immediately unsafe conditions, and should be reviewed
with site management.
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Hazardous Streams
A liquid stream may be considered hazardous because of its potential to flash or
vaporize and thereby create a fire or explosion hazard, or because it is toxic,
corrosive or otherwise unacceptable to allow it to be released to grade. Such liquids
should be routed to a containment system. This may be another process vessel, a
tank or a suitably designed flare system.
6.6.4
Two Phase Releases
Releases that may be combined vapor/liquid releases should be treated as liquid
releases and routed to the appropriate disposal location. These releases generally
originate from liquid filled vessels (boiling relief cases or flashing fluids) or from
vessels with small vapor spaces which could have liquid entrained in the relief
stream.
Generally, two phase releases are not safe to route directly to atmosphere and
should be routed to a closed system. If both the vapor and liquid phases are nonhazardous and the quantity of potential release is acceptable, grade disposal may
be considered. Any release of two phase material to a location other than a closed
system should be reviewed with site management.
6.6.5
Prevention of Liquid Releases
Occasionally, conditions are encountered where a liquid or liquid-bearing release will
occur only in very unlikely situations. If extensive systems are required to contain
the release, assuming the risk of the release to grade, atmosphere or into a flare
header designed primarily for vapor may be worth considering. Usually, the
conditions that make this viable are conditions in which the quantity of the release
would normally make it unsuitable for grade disposal, but the release in itself is not
immediately hazardous.
Some of the considerations that should be addressed during this assessment are:
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The potential hazard if the release does occur must be assessed. If an
immediately hazardous condition (such as release of volatile liquids to grade)
would be created, alternate disposal methods should be considered.
o
If the release occurs, the cause of the release must be readily correctable. If
the cause of the release cannot be rapidly and simply eliminated, alternate
disposal methods should be considered.
o
Additional controls and interlocks to minimize the chances of a liquid release
occurring should be added. These may take the form of a high level cut-off
or bypass or back-up control paths to either restrict fluid entering the system
or to significantly increase the flow of fluid leaving the system. These
systems do not eliminate the need for a properly sized pressure relief
system, but can substantially reduce the risk of a release occurring.
An example of this type of application is a feed surge or flash drum which may have
a non-volatile liquid release if the drum were to fill up with liquid. If all other potential
releases are vapors suitable for release to atmosphere, assuming the risk of a liquid
release may be acceptable if the above criteria are met.
This case also occurs in systems where a vapor stream is routed to a flare system
designed for vapors and the presence of liquid is not desired. An example is a flash
drum which has a relief stream containing H2S and which is routed to flare. If the
drum fills with liquid, liquid would be introduced into the flare system. While
releasing liquid into a flare system may not be desirable, it may be acceptable if
multiple failures would have to occur for it to happen.
6.6.6
Pressure Relief Device Failure
In evaluating the potential hazards associated with routing of a potentially liquid
stream to grade or sewer, the possibility of a pressure relief device failure occurring
should be considered. While not common, pressure relief valve springs can fail,
which would result in the valve opening during normal operation, perhaps requiring a
unit shutdown. If a stand-alone rupture disk is used, the possibility of the disk failing
prematurely should also be considered.
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DISPOSAL INTO A PROCESS
Routing of a pressure relief stream back into a process rather than to the environment or to
a disposal system may be the most economical and safest alternative for many streams.
This is particularly true of liquid or liquid bearing streams. However, there are several
considerations that must be addressed before it is decided to route a relief stream back into
a process.
6.7.1
Capacity
The ability for the receiving process to absorb the relief stream must be analyzed. If
the introduction of the relief stream can cause over pressure, then additional
pressure relief capacity on the receiving process may be required. Applications with
all vapor or sub-cooled liquid releases are fairly straightforward. If a tower or flash
drum is receiving a hot or two phase relief fluid, the fluid may flash or otherwise
interact with the fluid already in the receiving process and produce more vapor than
the normal exit paths can absorb.
6.7.2
Destination Pressure
The potential back pressure caused by the receiving process must be considered in
sizing pressure relief devices. Sometimes the back pressure in a process being
considered is too high for a reasonably sized pressure relief device and alternate
destinations have to be considered.
Usually, the fully accumulated relieving pressure of the receiving process should be
taken as the destination pressure. If any other pressure is taken to be the receiving
pressure, the system needs to be carefully reviewed to establish what the maximum
pressure of the destination system could be when the originating system is relieving.
6.7.3
Process Upsets
The impact of the release on the receiving process must be considered.
Introduction of the release stream may cause unacceptable upsets in the receiving
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process. A hot release may cause a sudden flash in the receiving process which
may either damage the equipment (i.e. lift fractionator trays) or cause the pressure
relief valves on the receiving process to lift. This may even become a sizing case
for the receiving process pressure relief system. Even if an immediate increase in
liquid or vapor loading does not occur, contamination or composition changes
caused by the release may be unacceptable.
The nature of the incident that causes the release and alternatives to process
disposal should be considered in this evaluation. It may be that a significant upset
to the receiving system is more acceptable than other alternatives.
6.7.4
In-Service Requirements
A very important constrain which should be evaluated is that the receiving process
must be in service at all times when the source of the release stream requires over
pressure protection. This generally means that the protected equipment must be
taken out of service and secured before the process that receives the relief stream
can be taken out of service. This type of application is usually found in pumped
circuits that require over pressure protection for blocked discharge or fire relief
cases. These systems usually discharge into a part of the process with lower
operating pressure, such as a flash drum or tower.
In-service requirements can place significant restrictions on operating flexibility if the
originating and receiving equipment are not closely tied together. Even if they are
parts of the same process, it may be required that the source equipment be secured
(usually drained and vented) instead of just being placed in a filled, standby mode.
6.8
CLOSED DISPOSAL SYSTEMS
Closed collection and disposal systems provide a means of preventing the release of
substances which cannot or should not be routed to atmosphere, grade or back into the
process. These systems provide a means of recovering release streams which have value
or safely disposing of those streams which cannot be economically recovered and reused.
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These collection systems may have sources from normal design releases from processes
(such as those from a process into a fuel gas system), emergency controlled releases from
pressure control and depressuring systems or emergency releases from pressure relief
devices.
The major disadvantage of closed systems is their cost, the need to provide enough
capacity for all potential release cases and the maintenance and operating attention they
require.
6.8.1
Intermediate Collection Systems
Intermediate collection systems are systems which collect vapor releases from
controlled sources such as pressure controlled process releases or small manual
releases from sources such as normal shutdown depressuring. Typical intermediate
collection systems are fuel gas systems or low pressure vapor recovery systems.
Releases into an intermediate collection system usually are treated in some manner
and recovered for use somewhere else in a facility. In refineries, the collected
vapors are usually used as fuel gas. If the collected vapors are in excess of needs,
the excess is usually released to a flare system for disposal.
It is not a usual practice to tie pressure relief valves into an intermediate collection
system. However, if the collection system is properly designed for the release, it
can be done. The primary requirement is that the intermediate system be equipped
with its own pressure relief system which is designed so that the collection system
pressure cannot rise high enough to adversely affect the function or capacity of
pressure relief devices discharging into it. Additionally, the system piping must be
sized to handle the flows from the pressure relief systems in addition to other flows
into it which may occur at the same time.
6.8.2
Flare Systems
Flare systems consist of a suitably designed combustion assembly, liquid knock-out
and pump away systems and the piping networks which connect them to the
pressure relief valves, control valves, block valves, etc. which discharge into the
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system. Flare systems may act as base-loaded disposal systems for vapors which
are in excess of refinery needs, and as disposal systems of emergency releases
from control valves, depressuring systems and pressure relief devices. They may
be fed directly from various processes or by an intermediate collection system or
both.
Refinery flare systems are primarily intended to contain and dispose of vapor
releases, but must also be capable of handling liquids which may be released to
them as a result of unit upsets, equipment failures, emergencies or mis-operations.
Liquids may also collect in the system due to condensation of heavy, hot vapors or
steam if they are released into the system. The flare itself is intended only to burn
vapors, so any liquids released into the flare header system must be separated from
the flare vapor and separately handled.
6.8.2.1
Low Temperature Fluids
Low temperature fluids require special consideration, particularly if there is
a possibility of low boiling liquids entering the disposal system.
Autorefrigeration will occur as liquid boils at the reduced pressure. If the
equilibrium temperature is sufficiently low, piping and drums fabricated of
materials designed for low temperature will be required to eliminate the risk
of brittle fracture. In such circumstances, consideration should be given
either to a completely separate low temperature system or isolation of the
stream until it reaches a knockout drum where the liquid can disengage.
6.8.2.2
Viscosity and Solidification
In the selection of a disposal system for liquids and condensable vapors,
the production of highly viscous or solid materials warrants consideration.
The design of a disposal system for such materials may require steam
tracing of valves and discharge lines. The formation of gums, polymers,
coke, or ice that might prevent safe operation of the discharge system
should also be considered in the design.
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Miscibility
Solubility or miscibility of the material with water and avoidance of the
formation of emulsions should be considered in the selection of a disposal
system.
6.8.3
Vapor Recovery
A vapor recovery system may be connected to a flare system. This design allows
for recovery of vapors which otherwise might be flared, but does not reduce the
sizing requirements for the flare system. Many times the flare gas is treated and
routed to the refinery fuel gas system. Collection and recovery of tankage boil-off
vapors may also be warranted.
6.8.4
Incinerators and Burn Pits
Disposal by an incinerator is based on combustion with induced air in a refractory
lined chamber. Vendors normally provide design.
For relief loads that contain large quantities of liquid and gas, disposal is a problem.
Before flares or incinerators can be used, the flow has to be separated into a liquid
and a vapor stream. A burn pit is suitable for mixed relief.
A burn pit is simply a shallow earth or concrete-surfaced pool enclosed by a dike,
with a liquid/vapor inlet pipe through the wall, and provided with pilots and igniter.
An incinerator should normally be considered for liquid disposal rather than a burnpit which is therefore not treated in detail.
The center of the flame is assumed to be 1.5 pool diameters from the center of the
pool, in the direction of the point where radiant heat density is being considered.
This assumption is used to allow for flame deflection by the wind.
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Note: This applies only to burn pits where a good distribution of the flames can be
guaranteed. A good distribution of flames for butane or lighter hydrocarbons, for the
large quantities often involved, can only be guaranteed by installing more burners.
6.8.5
Liquid Handling Systems
Liquid handling systems are designed when necessary to provide a means of
collecting and recovering liquid releases into flare headers, intermediate collection
systems or direct releases from processes. Usually, they consist of separation and
surge vessels, pump-out pumps or other liquid movement systems and the lines
connecting the surge vessels to the sources of release. Normally, the separator
vessels of liquid handling systems have their vapor spaces connected to
atmosphere, a flare or an intermediate vapor collection system.
There are two basic types of liquid handling systems. Those systems designed
specifically to handle all-liquid releases from liquid sources and those systems
designed to catch liquids which may be released to, or may condense in, a flare
header system.
The first type of liquid handling system may consist of piping directly connected to a
low pressure collection tank. The tank may be equipped with vents to allow any
vapor associated with the liquid release to be directly released to atmosphere or
released to a flare or vapor recovery system. This type of system is intended only
for non-volatile liquids which can be safely contained at atmospheric pressure.
The second type of liquid handling system is designed to handle a broad range of
liquid releases and is generally associated with a flare system. Depending upon the
expected liquid loading for the system, the system may have several forms:
o
A single separator/surge vessel, located near the flare. This design is
intended primarily for systems which have low potential liquid content.
o
A separator at each unit or groups of unit with an additional separator at the
flare. Depending upon the type and quantity of releases, this system may
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have all releases routed to a common header or have a separate liquid
header.
In a single separator system, the separator vessel must be large enough to catch all
potential liquid releases and still allow adequate free area for liquid/vapor
disengagement. If the vessel or pump-out pumps are not large enough, the vessel
may fill with liquid, which ends up going out of the flare.
In the local separator system, the primary liquid containment is at the units, with the
flare knock-out drum intended only to catch liquid which is not separated at the unit
separator vessels or which condenses in the line to the flare. The choice of a local
separator may simplify the design as some liquid entrainment may be acceptable at
the unit separator vessels and the piping between the units and the flare may not
have to be designed for possible liquid slugs. This simplification is at the cost of the
additional separators and their associated equipment, but often is the only effective
way of handling significant potential liquid releases.
6.8.6
Treating Systems
Some types of releases are not suitable for direct releases into a recovery or
disposal system and require some type of treatment near the release point.
Typically this treatment may be necessary because the fluids are too hot, contain
large amount of condensables, or are corrosive or reactive. The details of treatment
systems vary with the characteristics of the release, but typical refinery applications
are outlined below.
Some systems which receive hot condensable vapors are provided with systems
that cool the releases and recover any liquids that might otherwise condense in flare
lines. These systems typically use direct liquid quench or coolers and condensers.
Examples of such systems are coker blowdowns systems that handle both routine
coker drum blowdown gases and the hot releases from coke drum pressure relief
valves.
Liquid handling systems may also contain provisions for treating of the release
stream before it is processed further (neutralization, for example).
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DESIGN CONSIDERATIONS
API RP 521 (Third Edition), paragraph 5.4 provides extensive discussions of design
requirements for various types of emergency release collection, treatment and disposal
systems. An engineer should be familiar with this material before attempting to specify
disposal systems. Some of the key considerations are outlined below:
6.9.1
Atmospheric Releases
When vapor relief streams are routed to atmosphere, the release piping should be
designed to maximize the safety of release. Some design requirements are:
o
The release point must be located such that nearby personnel and
equipment are not endangered. This generally means that the release must
be at least 3 meters above any platform or ladder within an 8 meter radius of
the release point.
o
The release outlet piping must be sized for adequate dispersion. API RP
521 paragraph 4.3.2.2 provides some guidelines which are based upon the
exit Reynolds’ number and ambient conditions. This evaluation should be
performed at the minimum stable relief device flow. For spring opposed
pressure relief valves, this is 25% of capacity.
o
Potentially flammable release piping should be equipped with a snuffing
steam supply. The snuffing steam generally is not intended to extinguish a
tail pipe flame which might occur during full flow, but will allow any flames
from a leaking or improperly reseated pressure relief valve to be
extinguished.
o
6.9.2
In order to minimize fugitive emissions, rupture disks can be installed under
pressure relief valves which are vented to atmosphere. The need for such
installations should be established as part of each project scope.
Intermediate Collection Systems
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Intermediate vapor collection systems generally are intended to have limited
functions for handling of emergency releases, but the effect of emergencies must be
considered. Most intermediate collection systems have control valves connected to
them which will fully open under over pressure situations. Provided that pressure
control instrumentation is functioning, control systems will preferentially release
through pressure control valves before the pressure relief valves lift.
Therefore, the potential total emergency release into the system should be
quantified and the system should be equipped with suitably sized pressure
controlled release control valves and pressure relief valves which release into a flare
system (in some cases, pressure relief valve release to atmosphere may be
acceptable). The collection system must have its own pressure relief system, but
the size and set pressure for these systems will be affected by the back pressure
tolerance of the devices discharging into it. If pressure relief valves are discharging
into the system, the set pressure and capacity of the collection system pressure
relief valves must be such that the performance and capacity of any pressure relief
devices which discharge into the system are not adversely affected.
The potential for liquid being discharged into intermediate collection systems must
be considered. While most systems are designed only for vapor releases, liquid
could easily be discharged to them during an upset. An example is a pressure
controlled release from a distillation tower overhead accumulator. Normally this
release is vapor, but under a reflux failure case, the accumulator could easily fill with
liquid, which would flow into the collection system through a pressure control valve
which has opened to attempt to keep the column pressure from rising. In this
manner a large amount of liquid could enter the system and provisions for this
should be made in the collection system design.
6.9.3
Flare Systems
Design of flare systems consists of the following major activities, which are
presented in detail in Chapters 5, 7, 8 and 9:
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o
Identification and quantification of the simultaneous flow rates which may
enter the system.
o
Sizing of liquid knock-out drums and pump-out capacity.
o
Sizing and selection of the flare.
o
Sizing of flare headers and design of required supports.
6.9.3.1
Low Pressure Flares
Disposal of relief from a tankage area requires separate consideration. As
most relief valves in the area have low set points, this causes a problem for
disposal by flare. There is only a low pressure available at the valve outlet
and back-pressure is severely limited. As a result tankage areas will often
have their own low pressure flare.
Low pressure flare and relief systems are often used offshore. Stabilized
oil, fuel gas, injection water deaerators, closed drain systems, oil-water
separators, etc., operate close to atmospheric pressure and can withstand
only low back pressure from the relief system.
An economic optimization is required to determine whether a separate high
pressure flare system is justified as well.
6.9.4
Vapor Recovery
6.9.4.1
Flare Gas Recovery Systems
Figure 6.1 shows a conceptual design for a flare gas recovery system.
Typically, the system consists of one or more compressors whose suction
is directly connected to the flare header. The compressed gas is usually
routed to some type of treating system appropriate for the gas
composition, then to fuel gas or processing systems.
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Flare gases can have widely varying compositions which must be
evaluated during specification of recovery systems. The potential for
materials which are not compatible with the flare gas treating systems or
ultimate destinations must be determined. For example, streams
containing acid gases typically are routed directly to the flare, thereby
bypassing the recovery system.
o
Sizing
Flare systems are used for both normal process releases and
emergency releases. Flare gas recovery systems are seldom sized
for emergency flare loads. Usually, economics dictate that capacity
be provided for some “normal” flare rate, above which gas is flared.
Flare loads vary widely over time, and the “normal” rate may
represent some average flare load, or a frequently encountered
maximum load. Actual loads on these systems will vary widely and
they must be designed to operate over a wide range of dynamically
changing loads. Flare gas recovery systems often are installed to
comply with local regulatory limits on flare operation, and, therefore,
must be sized to conform to any such limits.
o
Flare Tie-In
A major consideration in flare recovery system design is preservation
of a path to the flare for emergency releases. The flare gas recovery
system must be designed as a side stream from the flare header.
Main flare flow should not be through any compressor knock out or
suction piping. The tie-in line to the flare gas recovery system should
come off the top of the flare line to minimize the possibility of liquid
entrance.
o
Positive Pressure Requirement
Some method of ensuring a positive pressure on the flare gas
recovery system must also be provided. Figure 6.2 shows some
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methods of doing this while preserving a reliable open path to the
flare.
a) Water Seal
The most positive and preferred method for preventing air ingress
from the flare to the blowdown system is the installation of a water
seal vessel between the flare knockout drum and the flare itself. The
seal provides a relatively constant low backpressure on the flare
header and provides a narrow, but usually adequate control range for
the flare gas recovery control system. The water seal should be
designed to function over the pressure for which the flare gas
recovery system is designed to operate. At higher release rates, flare
gas flows through the seal and to the flare. Design provisions must
be made to maintain the seal level, prevent high flare rates from
carrying the seal water up the flare and prevention of seal freeze-up.
See Chapter 8 for typical seal drum design considerations.
b)
Relief Valve Method
If process requirements are such that the narrow operating ranges
afforded by water seals cannot be accepted, an alternate method is
to use a fail open control valve to regulate the suction pressure of the
flare gas recovery system. A positive path to the flare is provided by
installing a low pressure, high capacity pilot operated pressure relief
valve around the control valve. The sensing line for the pressure
relief valve pilot shall be provided with a clean gas purge and a
backflow preventer.
The sizes of the control and pressure relief valves can become quite
large. The flare header system must also be studied to verify that the
backpressure imposed by the pressure relief device (assuming the
control valve is closed) at full header load will not induce
unacceptable back pressures on devices releasing into the headers
at the processing units.
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Other Methods
Alternatives to the use of a pressure relief valve are installation of
non-reclosing devices such as rupture disks or rupture pin devices.
These installations must also be carefully reviewed to ensure that the
devices operate at as low a pressure as possible, and that they do
not cause unacceptable back pressure.
If a control valve must be used in the flare line to regulate flare gas
recovery system suction pressure, the control valve should be of a fail
open design and be interlocked to go fully open upon a higher than
normal header pressure, high oxygen content, or when the
compressors are unloaded or shutdown. These interlocks are not a
substitute for a positive path around the control valve.
Provisions must be made to prevent back flow of air from the flare
into the flare gas recovery system. All compressors should be
equipped with highly reliable low suction pressure shutdown controls.
Consideration should also be given to installation of additional
instrumentation in the header section between the flare and the
compressor suction take-off to detect reverse flow and automatically
shut down the flare gas recovery system.
o
Location
Typically, flare gas recovery systems are located on the main flare
header downstream of all unit headers tie-ins and at a point where
header pressure does not vary substantially with load. Locations
upstream of process unit tie-ins should be carefully considered
because of the potential for back-flow and high oxygen
concentrations. Limited downstream tie-ins for material not suitable
for recovery may be required.
o
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Flare gas recovery systems must operate over wide ranges usually
within very narrow suction pressure bands. A typical system might
operate over a suction pressure range of 2-5 in (50-125 mm) of water
to 10-12 in (250-300 mm) of water. The flare gas recovery
compressors should be equipped with several stages of unloaders
and a compressor recycle valve. Suction pressure is normally
maintained via the recycle valve, with additional loading and
unloading of the compressors used when limits of valve opening or
closing or suction pressure are reached. Usually, the controls are set
up to sequentially load and unload the compressors.
The possibility of significant liquid in flare systems is usually quite
high. Liquid knock out vessels should be provided for the
compressors with automatic shutdown of the compressors on high
suction drum levels. Other mechanical protection systems may also
be required for the compressors. These systems may either shut
down or just unload the compressors.
6.9.5
Incinerators
Where a simple or less expensive form of disposal is not suitable, the relief products
can be burned in an incinerator. The most commonly used type for waste gas
incineration is a single hearth or firebox type. Firebox type incinerators consist of a
refractory lined rectangular combustion chamber. They typically have a pilot burner,
a rich fuel burner, and a waste burner. They normally operate with an induced air
supply so that a subatmospheric pressure exists in the firebox. An important factor
in the design of an incinerator is to achieve a stable flame. With an unstable flame
there is the possibility of unburned gas being discharged. Temperatures range from
1470 – 2550 °F (800 - 1400 °C). Incinerators often have a firetube boiler section
followed by a preheater and a scrubber column. The scrubber column is employed
for noxious or hazardous gases.
Other types of incinerators include:
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Multi-hearth
o
Fluid Bed
o
Flash Type
o
Rotary Kiln
o
Cyclone
o
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Liquid incinerators are designed to eliminate the problem of disposal of liquid waste
materials. They are readily adaptable to the disposal of virtually any liquid waste
combustible, including streams which contain inorganic or other non-combustible
materials present in combination with the oxidizable contaminants.
Where suitable storage is provided, it should be used to store the liquid relief
products and add them in a controlled manner to the incinerator to minimize large
firing rate changes.
Liquid waste streams which contain more than about 9900 Btu/lb (23000 kJ/kg,
5500 kcal/kg) units can usually be incinerated directly in a thermal burner without
the requirements of an auxiliary combustion chamber. In some instances, where
the content is less than 9900 Btu/lb (23000 kJ/kg, 5500 kcal/kg) it is possible to
enrich the liquid waste with fuel gas or other fuel.
6.9.6
Liquid Handling Systems
Liquid handling systems, whether intended for liquid release streams or for liquid
components of flare or intermediate collection systems, should have receiver
vessels sized for the expected liquid loading from the unit headers. Usually these
knock-out vessels are connected to a system which pumps away the collected
liquid, and are connected to a flare header which allows any vapors in the streams
to be released.
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Depending upon the use of the system and the presence of other knock-out vessels
on the vapor side, the vessels may have to be sized to handle all liquid coming to
them plus leave sufficient vapor space for efficient vapor/liquid disengagement.
This is particularly true for flare knock-out vessels. However in some cases it may
be justifiable to size local separator vessels for minimum vapor space and allow
some liquid carry over. The secondary vessel at the flare should then have
sufficient capacity to handle any liquid which may come from the primary separator
vessels or which may condense in the flare lines.
Pump away capacity is a primary consideration. The expected liquid flows and their
durations must be established in order to size the separator drum and the pump
away capacity. The receiver drum should be sized to handle the quantity of liquid
expected from a single release without the need for pump-away.
6.9.7
Treating Systems
6.9.7.1 High Temperature Relief Streams
Hot liquids and vapor may be cooled and condensed by one of the
following methods:
o
Direct Quench
Pressure relief valves that discharge hot condensable hydrocarbon
vapors or liquids may be piped into a separate header that terminates
in a quench drum. In this service, quench can reduce the
temperature of the relief stream and may permit the use of less
expensive materials in downstream equipment. Cooling also
condenses some of the less volatile components and can reduce or
prevent the release of hot condensable vapors to the atmosphere. A
quench drum is a vessel equipped to spray a quenching liquid down
through the hot discharged gases as they pass at reduced velocity
through the drum. The quenching fluid may be water, gas oil, or
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another suitable liquid. The used liquid collects in the bottom of the
drum for subsequent removal.
One type of quench drum is a vertical vessel containing baffles that is
connected by a means of a conical transition to a vent stack or flare.
The condensable hydrocarbon material is fed into the drum below the
baffles. Water is introduced into the drum above the baffles at a rate
that depends on the temperature and the amount of hydrocarbon
material being fed to the quench drum. The water, spilling over the
baffles, desuperheats and condenses the hydrocarbon vapor, knocks
out entrained hydrocarbon liquid, and cools down the hydrocarbon
liquid collected in the bottom of the drum. The uncondensed vapor
and any steam formed pass up the vent stack or enter a flare system
(see Figure 6.3).
o
Submerged Discharge
The submerged discharge system is a relieving system that
terminates in parallel laterals submerged in a water-filled sump.
Holes are cut in the bottom of the laterals throughout their length,
imparting downward flow to the discharged effluent to obtain
maximum agitation, cooling, and condensing. Provisions must be
made to maintain a liquid level in the sump while the blowdown
system is being used. The discharge is drained from this sump into a
separator, where the oil and condensed vapors are removed from the
water. The submerged discharge system is not extensively used in
present-day design.
o
Indirect Condensing or Cooling
The use of shell-and-tube heat exchangers or coil-in-box coolers has
the merit of separating cooled or condensed material immediately. In
addition, the coil-in-box cooler can remove some heat in emergencies
when no cooling water is flowing.
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Scrubbing Systems
When the relief vapors are objectionable and cannot be discharged to
atmosphere and when the products of combustion are also objectionable
and the relief cannot be flared, some form of treatment of the relief is
required.
The treatment process will be specific to the materials in the relief and is
designed as a small process unit. The most frequently used process is a
scrubber/absorption system where the toxic/corrosive material is removed
from the relief vapors, the remaining vapors are then vented or flared and
the bottom liquid either recycled or disposed of as a liquid.
It may be required to install fans to provide the necessary suction to draw
the gas through the absorption plant.
When investigating the process some general considerations are as
follows:
o
Chemical absorption is usually preferable to physical absorption due
to the larger quantities it can remove.
o
Co-current systems are simpler and create less pressure drop than
counter-current systems.
A typical scrubber system is shown in Figure 6.4. This is a chlorine
absorption system which is designed to absorb normal vents and
emergency releases from a chlorine plant to prevent chlorine emissions.
6.10
REFERENCES:
1.
Briggs, G.A., “Diffusion Estimation for Small Emissions”, Environmental Research
Laboratories 1973 Annual Report, National Oceanic and Administration Air
Resources Atmospheric Turbulence and Diffusion Laboratory (May, 1973).
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Hoehne, V.O., Luce R.G., and Miga, L.W. “The Effect of Velocity, Temperature, and
Gas Molecular Weight on Flammability Limits in Wind-Blown Jets of Hydrocarbon
Gases”, Report to the American Petroleum Institute, Battelle Memorial Institute,
Columbus, Ohio (1970).
3.
Loudon, D.E., “Requirements for Safe Discharge of Hydrocarbons to Atmosphere”,
API Proceedings, 43, 418-433 (1963).
4.
Taylor, J.F., Grimmett, H.L., and Comings, E.C., “Isothermal Free Jets of Air Mixing
with Air”, Chemical Engineering Progress, 47, 175-180 (1951).
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TABLE 6.1
TYPICAL THRESHOLD LIMIT VALUES FOR
TOXIC OR HAZARDOUS CHEMICALS
FOUND IN REFINERIES
TWA (1)
mg/m3
STEL (2)
mg/m3
Carbon Dioxide
9,000
18,000
Carbon Monoxide
55
440
Chlorine
3
9
Hydrogen Sulphide
15
27
Nitrogen Dioxide
9
-
Sulphur Dioxide
13
-
Notes:
(1) Threshold Limit Value - Time weighted Average (TLA - TWA) the time weighed average
concentration for a normal 8 hour workday or 40 hour working week, to which nearly all workers
may be repeatedly exposed, day after day, without adverse effect.
(2) Threshold Limit Value - Short Term Exposure Limit (TLV - STEL) the maximum concentration to
which workers can be exposed for a period of up to 15 minutes.
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TYPICAL FLARE GAS RECOVERY SYSTEM
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FIGURE 6.2
FLARE GAS RECOVERY
INLET PRESSURE CONTROL SYSTEM
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TYPICAL QUENCH DRUM
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TYPICAL SCRUBBER SYSTEM
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7.0
RELIEF SYSTEM PIPING AND SEALING/PURGING
7.1
DESIGN CONSIDERATIONS
7.1.1
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Piping Layout Guidelines
Detailed design guidelines for piping layout are presented in the client’s engineering
standards or Fluor Daniel in-house piping layout guidelines. The process engineers
must be aware of them when piping isometric drawings are reviewed. The typical
guidelines are summarized below.
Pressure relief valves are generally located directly on the equipment or piping
being protected. Relief valves can be located on the piping from equipment as long
as inlet loss and other criteria are acceptable. Pressure relief valves which
discharge directly to the atmosphere are equipped with a minimum ½ inch weep
hole at the piping low point. Snuffing steam piped to the discharge line is required
for the relief valves relieving combustible gas. Typical relief valve installation
configurations are provided in Figures 7.5 through 7.11.
The following guidelines apply to relief valves which discharge to a collection header
system:
o
Top entry into the header is preferred over side or angle entry. Entering at a
45° angle in the direction of main flow is common for the relief valve
discharge pipes, however, it is not mandatory.
o
Pressure relief valves are normally located at a higher elevation than the
header to provide drainage.
o
Discharge piping from relief devices which are located below the header is to
be arranged to rise continuously to the header entry point. A drain
discharging to a safe location is required at the piping low point.
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Relief system headers that may be in contact with liquid are to be sloped to
knockout drums. The minimum slope is 8.3 in/330 ft (21 cm/100 m) of piping length
taking into account piping deflections between supports. Traps or other devices
with operating mechanisms cannot be used for liquid knockout service. (Refer to
Chapter 8.0 for design of knockout drums).
When it is absolutely necessary to install a riser in the flare header, as when
crossing a roadway, provisions must be made for draining the header ahead of the
riser.
Other relief valve installation criteria are provided in Chapter 3.
7.1.2
Design Temperature
The design maximum temperature to be used for analyzing stress in the relief
system piping is the highest temperature achieved downstream of the relief valves.
This information is available from the individual relief valve summaries. Generally,
fire relief cases establish this design temperature for laterals from the relief valve to
the subheader. Fire relief case temperatures are normally not considered in
establishing pipe wall thickness.
The design temperature selected above may be reduced in the downstream piping
by calculating heat loss to be ambient air. This may be performed in several zones
to reduce the need for large expansion joints associated with elevated temperatures.
This credit is to be applied to pipe stress only and not to hydraulic calculations.
Where the relief system is divided into several or multiple collection headers (such
as low and high pressure headers), several or multiple design temperatures may be
assigned; one per header.
Design minimum temperature is established based on the autorefrigeration
temperature of relieving fluids.
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Design Pressure
The minimum recommended design pressure for piping is 50 psig (3.4 barg, 3.5
kg/cm2 G) to correspond to the K.O. drum design pressure. This minimum is to be
raised, if necessary, to provide a margin of 10% over the maximum pressure
calculated for the design relieving case. Due to the significant pressure drop
associated with relief valve discharge laterals, a design pressure for the main relief
header lower than the lateral is acceptable where significant cost savings are
anticipated. Flare header systems are generally proven by weld x-rays and/or by
air/nitrogen testing. Hydrotesting is rarely used and is not recommended for large
header systems.
7.1.4
Stress
Relief system headers are subject to a wide range of operating conditions and
subsequent shock. Thermal stresses may result from entry of cold or hot materials.
Shock loading may result either from sudden releases of compressible fluids into the
multi-directional piping system or from impact action of liquid slugs at points of
changes in direction. The following process data is required by the piping stress
engineer to ensure proper design of the piping system:
7.1.5
o
Probable combinations of relieving conditions to be handled by the piping.
o
Probable mass rate, density, and sources of any liquid slugs.
o
Location of relief devices discharging large loads.
Isolation Valves
Block valves are normally not installed on the inlet or outlet of relief devices.
Exceptions may include applications with dual (spared) PSVs on one vessel, a spare
(standby) vessel with a PSV, a PSV in parallel with an autodepressuring valve, and
thermal relief valves on piping. When block valves are requested by the client, they
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must be locked or car sealed open when the equipment or piping being protected is
in service.
The client may request that provisions be made to permit isolation of individual
sections of the relief header system. These provisions are to consist of a block
valve and a spectacle blind for each section being isolated. See Figure 7.4 for
typical block valve locations. The block valve must be locked open when the unit or
units in which the section of header is located are in service. The blind is installed
after the valve is closed.
Gate valves used as block valves must be installed in the horizontal or upside down
position. This approach is needed to avoid flow blockage if the disk breaks from the
stem.
No such precautions are required for a butterfly valve. Select a butterfly valve which
is designed such that breakage of the disk from the stem does not block flow.
7.1.6
Design Criteria for Relief Valve Inlet Piping
The relief valve inlet and outlet piping must meet the criteria listed in Table 7.1 with
respect to pressure drop, flow rate and relief valve type.
The flow rate to be used for the line sizing depends on the type of valve and may
vary from inlet to outlet lines for the same service. The “required flow” is the
maximum flow required by the process to prevent overpressure. Where an exact
match of the required relief valve orifice to that available cannot be found, the next
larger available size is selected. The larger orifice will pass more than that required
by an amount proportional to the orifice area ratio.
W rated = W req’d (Arated / Areq’d )
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Where:
7.1.7
Arated
=
Area of the selected valve orifice, in2 (cm2 )
Areq’d
=
W req’d
=
Area calculated as required to relieve the maximum process
2
2
flow to prevent system overpressure, in (cm )
Maximum process flow required to prevent system
overpressure, lb/h (kg/h)
W rated
=
Actual flow through the relief valve at set pressure plus
required accumulation, lb/h (kg/h)
Design Criteria for Relief Headers
o
Backpressure
The basic criteria for sizing the relief header is that the back pressure, which
may exist or be developed at any point in the system, does not reduce the
relieving capacity of any of the pressure-relieving devices below the amount
required to protect the corresponding vessels from overpressure. Thus, the
effect of superimposed or built-up back pressure on the operating
characteristics of the valves should be carefully examined. The discharge
piping system should be designed so that the built-up back pressure caused
by the flow through the valve under consideration does not reduce the
capacity of any pressure relief valve that may be relieving simultaneously.
Refer to Table 7.1, Section B, for the basis for these calculations.
Normal header operating pressure with minimal flaring depends primarily on
the backpressure imposed by the flare liquid seal or other flare system
controls. A typical liquid seal with 1 ft (0.3 m) seal depth imposes a
2
backpressure of about 0.4 psi (0.03 bar, 0.03 kg/cm ). If flare vapor
recovery is installed, system operability requires a larger liquid seal depth of
about 3 ft (1m), resulting in a normal backpressure of about 1.4 psi (0.1 bar,
0.1 kg/cm2). This level of backpressure generally has an insignificant impact
on flare system design. However, addition of an enclosed ground flare to the
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system may require the need to impose more significant flare header
backpressure to allow practical staging control. This should be carefully
considered in the overall system design.
o
Worst-Case Cumulative Required Capacity
When discharge manifolds and relief headers are sized, the relief
contingency that produces the greatest back pressure should be identified.
Any single relief contingency may involve several pressure relief devices.
Typical relief contingencies that may be considered include cooling water
failure, power failure, and fire.
In addition to the back pressure criteria, the determination of the flow rate to
be considered forms the basis for discharge line sizing.
•
Laterals and tailpipes from individual devices are sized based on the
capacity required or rated, as indicated in Table 7.1. It is anticipated that
the flow rates for sizing the inlet and outlet piping for the same relief
device may not be consistent.
•
Common header systems and manifolds in multiple device installations
are generally sized based on the worst-case cumulative required
capacities of all devices that may reasonably be expected to discharge
simultaneously in a single overpressure event.
Good engineering judgment should be applied to select the flow basis most
appropriate to each case. (See the ASME Code, Section VIII, Division I,
Appendix M-8).
o
Vapor Depressuring Systems
In designing vapor depressuring systems, precise pressure drop calculations
are usually not necessary. However, the built-up backpressure should be
limited based on the following considerations:
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a)
The ratings of fittings exposed to back pressure should not be
exceeded.
b)
Back-flow from the header into any connected process should be
avoided.
Piping Metallurgy
All inlet and outlet piping to and from safety/relief and depressuring valves are to be
fabricated from normal carbon steel for the service temperature above -20 °F (-29
°C), unless otherwise specified. This selection of carbon steel is to minimize piping
costs. For some fluids carbon steel is totally unsuitable if corrosion reduces the life
of the line significantly. A separate relief header may be required to handle very
corrosive fluids. Acid gases or sour gases (CO2, H2S, HCN) would be candidates
for special metallurgy. Schemes for warming up relief streams with steam or by
mixtures of warm gas shall not be taken into account for the material selection.
For minimum service temperatures below -20 °F (-29 °C) but above -51 °F (-46 °C),
the flare headers may be constructed of fine grain steel (low temperature carbon
steel, ASTM A333, Gr. 6 or equivalent) or normalized carbon steel without detailed
calculation of the stresses. The material selection of the piping downstream of
safety/relief valves shall be based on material having sufficient impact values for
local stresses over 50 Newtons/mm2.
For service temperatures below -51 °F (-46 °C), alloy piping must be used. Cold
discharges result not only from cold processes, but also from high pressure ambient
temperature liquids which auto-refrigerate due to flashing as the pressure is
reduced, such as liquid propane. Cold discharges will warm up as the gas flows
along the header either by gaining heat from the pipe and surroundings or from
mixing with warm gases, making carbon steel suitable downstream. It is not
unusual to use alloy steel for the local pressure relief valve discharge line and
carbon steel for the header it ties into.
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Where headers of different material have to be connected, the piping material with
higher impact value shall be used for a length equal to 50 pipe diameters with a
maximum of 33 feet (10 meters) upstream of the transition point of process
conditions, because of possible back-flow.
Relief system piping shall be designed to meet the full requirements of the piping
class that has been specified.
7.1.9
Winterization, Safety Insulation and Steam Tracing
Reliable heat tracing (electrical or steam) is required on the inlet and/or outlet
piping for relief valves if the discharged fluid is subject to freezing due to severe cold
weather or from the effects of flash cooling.
A frequent problem in cold climates is the formation of ice on the inside bottom of
the relief header due to heat loss to the atmosphere when there is little gas flow in
the relief header. Both relief valves and block valves will tend to leak slightly over a
period of time. A leak from a cool wet service could result in a gradual build-up of
ice immediately downstream of the relief valve and in the relief header with serious
damage resulting.
In polymer or similar service where the fluid solidifies or becomes extremely viscous
as it cools, the relief valve inlet line should be traced and treated as any other dead
leg in that service to prevent any restriction of relief flow.
Piping may become very hot during relief episodes and may require insulation for
personnel protection near operating platforms and walkways.
If freezing is possible due to water, tars, heavy oils, or other condensates being
present, provide heat tracing using the client’s standards or Fluor Daniel in-house
standards.
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LINE SIZING
7.2.1
Relief Valve Inlet/Outlet Piping Sizing
7.2.1.1
Introduction
These guidelines are intended to provide a standardized approach for
sizing relief valve inlet and outlet lines. Guidance is also provided on how
to resolve situations where the guidelines are difficult to meet.
Process engineer is responsible for sizing relief valve inlet and outlet lines.
Improper sizing of these lines can result in valve damage and/or
inadequate relieving capacity. Since system safety must be assured, it is
important that calculations be standardized and the key guidelines clearly
understood.
7.2.1.2
Basis
Inlet frictional loss and outlet back pressure must be checked for each
valve based on guidelines provided in API RP 520 and 521. An
interpretation of these criteria for the various valve types and different flow
regimes has been prepared and included as Table 7.1. This interpretation
has been reviewed and concurred with by several major relief valve
manufacturers.
High inlet frictional loss can result in valve chattering. Conventional valves
relieving vapor flow tend to pop open and relieve at near maximum flow
initially. If frictional loss is high at these flows, inlet pressure can be
reduced far enough to cause the valve to rapidly reseat. Under this
circumstance, in addition to the possibility of damaging the valve, the flow
which can be relieved can be severely limited, resulting in a safety hazard.
API guidelines limit inlet frictional losses to 3% of set pressure. Every
effort to meet these criteria should be made for new installations.
However, for existing installations or for new installations with certain
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difficult valve sizes, slightly higher losses can be considered with the
approval of the valve manufacturer and the client. Typically, inlet losses of
up to 4% may be acceptable to the manufacturer and the client, based on
past experience. One industry user reported accepting inlet losses up to
5% based on special valve adjustments, but required actual performance
tests before accepting the valves.
Excessive outlet back pressure can also severely limit the capacity of a
conventional relief valve by causing it to close. This becomes particularly
important for a valve discharging to a relief header, because back pressure
is impacted by the simultaneous releases from other sources. In addition,
variable back pressure introduced by other sources already venting to a
flare header can cause a valve to lift at a higher pressure than required.
Back pressure against a conventional relief valve should be limited to 10%
of set pressure per API guidelines, whereas balanced bellows and pilot
operated relief valves are much more tolerant of back pressure. As with
inlet losses, it is sometimes possible to obtain agreement with the
manufacturer and the client that a somewhat higher back pressure be
acceptable for a particular conventional valve installation.
The potential for a serious problem to develop due to inadequate hydraulic
design of relief valve piping is difficult to assess on an absolute basis
because relief valve manufacturers have done very few tests outside the
limits set out in the API guidelines. While manufacturers assure that their
relief valves will operate properly when the guidelines are followed and
recognize that there is some conservatism built into the guidelines, they do
not normally have sufficient data (or any incentive) to accurately predict
how far the guidelines can be exceeded without significant consequences.
The following factors tend to aggravate the problems resulting from
hydraulic restrictions:
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o
High pressure relief
o
Large relief valve size
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o
Small margin between required and rated valve capacity
o
Inability to isolate a leaking relief valve for repair
Responsibility for installations not specifically meeting the API hydraulic
criteria must clearly be accepted by the client and/or the valve
manufacturer.
7.2.1.3
Required Data
Hydraulic calculations should be completed for the controlling relief case
(basis for valve sizing). In addition, other significant relief cases which
involve different fluid phases should also be evaluated. For example, if a
liquid relief case requires 80% of the relief valve orifice area of the
controlling vapor case, both of these cases should be evaluated
hydraulically. However, if the required liquid relief case area is only 5% of
that of the vapor case, there is no need to evaluate that liquid relief case
hydraulically.
Once the case(s) to be evaluated are determined, the data required to
complete the hydraulic calculations is obtained as follows:
1.
Relief valve type and size from relief valve data sheet.
2.
Determine if required or rated flow should be used based on Table
7.1.
3.
Obtain required or rated flow from relief valve data sheet. Verify that
values shown represent flow from the relief valve being evaluated (if
multiple valves are required) at the correct over-pressure condition.
4.
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Estimated outlet back pressure provided on relief valve data sheet
(by Process).
6.
Verify pipe schedules based on line class and size.
Any missing data or uncertainties in the values listed on the relief valve
data sheet should be resolved before proceeding with the hydraulic
calculations.
7.2.1.4
Inlet Calculations
Once a relief valve lifts and fluid begins to flow, both the upstream static
pressure and the kinetic energy of the fluid act to keep the valve open. As
a result, only the frictional loss needs to be considered in evaluating inlet
pressure loss. Pressure reduction due to velocity head increases across
reducers should not be counted as an inlet loss. Unless changes in static
head occur during relief, static head can also be neglected in evaluating
inlet loss. Note that the nominal inlet line size must always be at least
equal to the pressure relief valve inlet flange.
As shown in Table 7.2, the ratio of relief valve orifice area to inlet line flow
area varies greatly for different relief valve sizes, ranging from 0.054 for a
1½D2 valve up to 0.554 for a 6R8 valve. Clearly, the larger this ratio, the
more difficult it is to satisfy inlet loss criteria. Table 7.3 presents the
maximum equivalent length of inlet line (size same as valve inlet) which
can be used while satisfying the 3% loss criteria for each size valve at
different operating pressures. This table can be used as a quick method of
selecting a preliminary inlet line size for the calculation and determining
how detailed a calculation may be required. For example, with almost any
reasonable piping configuration a 1½ inch inlet line will be adequate with a
1½D2 relief valve. In contrast, it is unlikely that a 2 inch inlet line will be
adequate with a 2J3 relief valve with most configurations. Also, the inlet
calculation can be simplified by using conservative assumptions for a
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1½D2 relief valve, whereas a more detailed and careful calculation is
probably required with a 2J3 relief valve.
Where it is likely that the selected inlet size will work easily, conservative
overall line lengths can be calculated to provide quick documentation of
system adequacy. However, in more difficult cases, it may be necessary
to carefully account for fitting lengths and include reducer angles, as
discussed in Section 7.2.1.6, in order to verify that a configuration is
acceptable.
If the initial inlet configuration results in excessive frictional loss, the
following options can be considered as a next step:
1.
Increase inlet line size.
2.
Change valve selection for lower orifice/inlet area ratio valve(s) which
still meet relief load requirements.
3.
Work with the piping designer to reduce inlet line length or number of
fittings.
4.
Change to pilot operated valve with remote sensing.
5.
Change to modulating pilot valve to reduce flow used for calculation.
6.
Obtain vendor agreement to exceed 3% guideline.
Option 6 should be used very selectively with new installations, as a great
deal of effort can be required to obtain written agreements. By locating a
pilot valve sensing line on the vessel or other protected equipment item,
the inlet loss only applies to the pilot line. Many pilot valves do not induce
flow through the pilot line, so inlet losses are essentially zero. Pilot lines
which do flow under relieving conditions must be sized to meet the 3%
criteria (normally these lines are generously sized for support purposes, so
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line loss is often negligible). If the pilot line tap is placed in a flowing line
downstream of the protected equipment, a pilot type connection should be
made which recovers the dynamic energy of the flowing stream. Flow
capacity of the pilot operated relief valve still must be reduced to account
for high inlet pressure drop and resulting lower inlet pressure.
7.2.1.5
Outlet Calculations
Unlike inlet loss criteria, it is the actual back pressure at the relief valve
outlet on which the API criteria is based. Outlet calculations are completed
by starting at the first downstream point where pressure is known and
calculating the pressure profile backwards up the line to the relief valve
outlet. This known pressure point is most frequently at the flare, the unit
relief header, or atmosphere (for a valve which vents to atmosphere). Trial
and error calculations may be required in the case of flashing liquid relief to
obtain the vapor/liquid ratio in a specific section of line.
Frequently, the flow at the first expansion upon exiting the relief valve
approaches sonic velocity. Sonic velocity is acceptable for emergency
relief conditions. However, sustained flow at sonic conditions can lead to
vibration induced piping failures and must be carefully considered in piping
design or avoided. Even if pressure drop is low in the downstream piping,
flow may be choked at this point and pressure fixed based on sonic
velocity. Whenever pressure changes are large across reducers with two
phase flow present, manual calculations should be performed to better
estimate the actual pressure change, if the impact on back pressure is
critical. Otherwise the upstream pressure of the expansion can
conservatively be assumed to be equal to the downstream side. If the two
phase stream approaches sonic velocity near the valve outlet, the
upstream pressure of the expander must also be checked to determine if
the stream is choked at a higher pressure than that calculated by
hydraulics. If the results are critical, it is best to review the calculations
with the valve manufacturer, since detailed calculations for sonic velocity
are beyond the scope of most normal hydraulic checks.
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Outlet flow from certain valve sizes under particular relief conditions can be
choked at the valve outlet at a pressure exceeding the 10% criteria for
conventional valves. Under these circumstances, the vendor and the client
must agree that the back pressure is acceptable or a different valve type
more tolerant of back pressure must be used. Table 7.4 shows the ratio of
orifice area to valve outlet area for standard relief valve sizes. As
suggested by the table, the 6R8 and 8T10 valves are particularly
susceptible to this type of behavior due to their high orifice/outlet area
ratios.
If the outlet configuration results in a calculated back pressure in excess of
that allowable by Table 7.1, then the following options can be considered
as a next step. The order of consideration will depend upon the nature of
the job. Design of new piping will be approached differently from a review
of existing piping.
1.
Increase outlet line size.
2.
Change valve selection for lower orifice/outlet area ratio valve(s)
which still meet relief load requirements.
3.
Work with Piping to reduce outlet line length and/or number of
fittings.
4.
Work with the Instrument Engineer to determine if back pressure can
be increased for balanced bellows or pilot operated PSVs.
Change valve selection to a type more tolerant of back pressure.
5.
6.
Obtain client and vendor agreement to exceed allowable back
pressure (conventional valves).
In calculating back pressure in a relief header system, the flow criteria
obtained from Table 7.1 for the case under consideration only applies to
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the immediate outlet and local laterals from the relief valve. Once the line
branches out to the larger relief subheader, the flow rate is assumed to be
the required relief flow rate, even if the local valve outlet line is being
evaluated at maximum relief valve capacity. Judgment is sometimes
required to determine where this break occurs between local valve outlet
piping and the relief header. Both increased diameter and distance from
the valve provide the volume and velocity damping which tends to smooth
out short term high flow rates from the relief valve.
7.2.1.6
Calculation Details
In cases where detailed calculations must be done to assure that a
particular configuration is satisfactory, a detailed isometric drawing or
sketch of the piping is required so that accurate line lengths between
fittings can be obtained. Several key features of a detailed calculation are
discussed in the following paragraphs.
Vessel outlet nozzle lengths must be included as part of the inlet line
length. Typical vessel nozzle lengths are provided in Table 7.5. Vessel
outlets can be simply modeled as 180 degree expansions from the
diameter of a vertical vessel (½ the diameter of a horizontal vessel) to the
diameter of the nozzle.
Branch tee connections are common in relief valve piping. A normally
conservative method of calculating pressure loss in a branch tee is to
calculate it as a branch tee at the larger diameter and corresponding flow
rate, followed by a reducer to the smaller line size and the flow rate
corresponding with the smaller line.
Pressure losses across reducers can be significant, particularly for inlet
calculations. While reducers are not standardized, it is reasonable to
include approximate angles for reducers where “normal” reducers are
being used. This refines the calculation to allow for lower pressure drop
than the standard calculation which assumes an immediate 180 degree
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expansion. Table 7.6 provides a list of typical reducer angles. Reducer
angles should only be used when it is certain that “typical” reducers are
being used. This excludes, for example, reducers used to simulate branch
tees, stub-in connections, field fabricated reducers (unless the angles can
be verified), and small pipe socket weld connections.
Significant pressure loss can occur in the piping section immediately
upstream or downstream of the relief valve, since this section must match
the size of the relief valve inlet or outlet. As a result it is frequently
desirable to provide a reducer as close as possible to the relief valve. In
that case the minimum straight length of “pipe” adjacent to the relief valve
is equal to the length of the weld neck flange. Table 7.7 provides a list of
typical weld neck flange lengths.
7.2.1.7
Follow-Up
After an acceptable piping configuration has been determined, all
necessary documentation must be updated to reflect any changes which
have been made. These particularly include updates of the relief valve
data sheet in conjunction with P&ID markups, and Process approval of the
final piping isometric drawings. Controls should be implemented on each
project to assure that Process approves any revisions to these approved
relief valve piping isometric drawings prior to their issue.
7.2.2
Line Sizing of the Main Relief Header
Where the ratios of pressure relief valve set pressures in a system are in the order
of 5 to 1 or higher, the feasibility and economy of separate high and low pressure
collection headers terminating in either a common knockout drum and flare or
separate knockout drums and flares should be investigated. This often results in
smaller headers (lower total cross sectional area).
When the maximum vapor relieving requirement has been established and the
maximum allowable header back pressure has been defined (as determined by the
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type of valves in the system and the applicable code requirements), the selection of
line size is then reduced to fluid flow calculations. Various proprietary methods can
be used by the responsible contractor to calculate the size of discharge piping when
the flow conditions are known. These range from treating the flow as isothermal,
with appropriate allowances for kinetic energy effects, to the more rigorous solutions
afforded by the adiabatic approach.
However, the isothermal equations are generally used for flare system evaluation,
since this approach gives the most conservative results. Either commercially
available programs (such as INPLANT and Flarenet) or Fluor Daniel’s in-house
FSAP program can be used for flare system hydraulic evaluation.
There is no definite velocity limitations which should be used for sizing flare main
and sub headers. The flare headers are sized based on hydraulic evaluations, to
meet backpressure limitations for all the relief valves in a system. Generally, flare
header velocities fall under 0.5 mach in order to meet the relief valve backpressure
requirements. Process engineers may select initial header sizes based on a velocity
of about 0.3 mach as a starting point before completing hydraulic calculations.
However, it is not unusual to have sonic velocity in the branch lines. Gas velocity
often reaches sonic at either expansions or header connections in the relief valve
discharge lines. Sonic velocity in the flare branch line is acceptable, if this condition
lasts only for a short time. The flare branch lines should be properly supported to
withstand the potential vibration due to the sonic exit. Sometimes clients require
increased line sizes to avoid sonic velocity at the header inlets.
The primary calculation will be based on the relief case that has the highest total
flare load as determined by the analysis described in Chapter 5. This same case
will also set the location and spacing requirements for the flare and flare knockout
drum for thermal radiation safety. The calculation will determine the back pressure
at each relief valve and verify that the back pressure limits have not been exceeded.
Additional pressure drop calculations must be made for all relief cases that have
higher flare loads than the primary case in any of the subheaders or laterals. These
cases, having greater branch line loads but lower total process loads than the
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primary case, may set the size requirement for many portions of the process unit
flare piping. These pressure drop calculations may involve a complete analysis
starting from the flare tip. Alternatively it is satisfactory to simplify the calculations
by starting with the pressure profile developed for the primary case and revising only
those upstream flare line portions having a higher rate. These abbreviated
calculations give higher pressures than complete calculations, but are acceptable if
they demonstrate that pressure relief valve back pressure limits are met.
7.2.2.1
Preliminary Sizing
Preliminary sizing of the headers can be performed by using the DarcyWeisbach or Moody formula with the restrictions given in the Crane
Technical Paper No 410 or 410M (See page 3-3 of 410M, 5th Printing,
1986). The restrictions are that the pressure drop for the segment must
not be more than 10% of the inlet pressure using inlet or outlet density.
For pressure drops greater than 10% and less than 40% use the average
density. For pressure drops greater than 40% divide the segment into
several smaller segments. These calculations are to be confirmed as the
design progresses with calculations more representative of isothermal or
adiabatic compressible flow.
7.2.2.2
API Method
An isothermal flow equation is presented in API RP 520, 6th Edition, 1992,
Appendix F and API RP 521, 3rd Edition, 1990, Appendix E, for sizing
relief manifolds. Where the outlet pressure, line size, equivalent line
length, Moody friction factor (see Table 7.8), and gas temperature can be
estimated, the inlet pressure to the header section being sized can be
calculated by trial and error. This is a good way to check computer
calculations for reasonable values.
M2 = 1.702x10-5 (W/P2 D2 )(ZT/mw)0.5
-7
2
0.5
M2 = 3.225x10 (W/P2 D )(ZT/mw)
M2 = 3.287x10-7 (W/P2 D2 )(ZT/mw)0.5
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W critical = 5.875x104 (P2 D2 )(mw/ZT)0.5
W critical = 3.101x106 (P2 D2 )(mw/ZT)0.5
6
2
W critical = 3.042x10 (P2 D )(mw/ZT)
2
7.3 (English)
7.3 (Metric)
7.3 (Metric)
0.5
2
f Le/D = (1/M22 )(P1 /P2) [1 - (P2 / P1 ) ] - ln (P1 / P2 )
2
7.4
Where:
f
=
Darcy-Weisbach or Moody friction factor
Le
=
Equivalent length of pipe, ft (m)
D
=
Pipe inside diameter, ft (m)
M2
=
Mach number at pipe outlet
P1
=
Pressure at pipe inlet, psia (bara, kg/cm2 A)
P2
=
Pressure at pipe outlet, psia (bara, kg/cm A)
W
=
Gas flow rate, lb/h (kg/h)
W critical =
2
Maximum mass flow at outlet conditions, lb/h (kg/h)
This mass rate cannot be exceeded at the outlet
conditions (temperature and pressure) of the pipe
segment because it would cause sonic velocity to occur at
that point. If the required flow is higher, then the line size
must be increased.
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Z
=
Gas compressibility factor
T
=
Gas temperature, °R (°K)
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Gas molecular weight
Flashing Liquid Headers
If the liquid content is small (less than one percent by volume) assume all
the liquid is vaporized and use one of the preceding methods for sizing
vapor relief headers. This will give a conservative design.
For systems with greater liquid content, use an approved two phase sizing
methods.
7.2.2.4
Depressuring Lines
Size depressuring lines by using the Fluor Daniel in-house calculation
spreadsheets. These calculations can be checked by using the method
developed by Lapple as described in section 4.8.5.
Where depressuring lines have been connected to the flare header, the
results of flow in these headers must be accounted for in the back
pressure calculations. If the potential flows through the depressuring
devices are larger than the relief devices on the equipment protected, then
consideration needs to be given to substituting the depressuring flows for
the relief flows on a case by case basis. Where the size is 12” or under
the velocity may be allowed to approach sonic. For larger sizes reduce
velocity to 0.8 mach or less. These limitations have been established to
minimize vibration and noise. Where the velocity is limited to 0.8 mach,
perform hydraulic calculations with the outlet pressure at the low or no flow
pressure in the relief mains (worst case for velocity).
While it is possible that all depressuring devices could be activated
simultaneously in an emergency, it is improbable. Where a clearly defined
administrative procedure has been developed and all operators trained in
the application of the procedure, it is anticipated that the maximum
depressuring rates will be something less than total. This quantity must be
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determined and used not only for the sizing of depressuring laterals but
also for calculating the total flows through the mains resulting from a major
upset condition.
7.2.2.5
Flow Basis in Combined Laterals
Where several relief loads are expected to be simultaneous, size the
combined lateral for the sum of the required flows.
7.3
COMPUTER MODELING OF FLARE HEADERS
Many hydraulic programs have been developed and used specifically for flare system
hydraulic evaluation. The following are some of the latest programs that are available in the
industry or in-house:
•
FSAP
-
In-house developed by Houston office based on PAL isothermal.
•
Flarenet
-
Hyprotech, been used in Camberley
•
INPLANT -
Simulation Sciences, Inc., used in Irvine office and proved to work
effectively
These hydraulic programs undergo continuous development and revision, therefore, lead
process engineers should select a suitable hydraulic program for the flare system
evaluation, based on project requirements, cost and client preference.
For complex flare system analysis, the latest version of the INPLANT program, has been
used to model hydraulic performance of the flare headers, subheaders and relief device
laterals. It is recommended that the network method described below be used, if INPLANT
is selected. Where velocities approach sonic, the portion of the header upstream of that
point must be evaluated separately. An example of the application of INPLANT for analysis
of an existing flare header is provided in Appendix B-8. The following discussions are
based on the INPLANT program.
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Network Method
The network method is used for any complex hydraulic network system.
Compositional data is not required to use the network method. Molecular weight or
density alone can be used to define stream data. Flow rates must be provided in
volumetric units. The network method is not capable of handling critical flow. If
critical flow occurs, the network method does not reach a converged solution.
However, the network method allows analysis of two or more flares (sinks) in a
single system as well as looped configurations where more than one path is
available for flow.
7.3.2
Flare Method
The flare method is specifically configured for flare systems. It contains special
capabilities in handling critical flows. Flow rate can be defined in either mass or
volumetric units. The flare method can only be used by defining, by component, the
composition of each relief stream. This method allows only one flare (sink).
7.3.3
Contingency Allowance
A design allowance for pressure drop calculations may be required to allow
contingency for fouling of the inside of the flare headers during service. This
contingency is to be considered for each project and an appropriate value
established. For the majority of systems no contingency is required since the
corrosive effects of oxygen are excluded and heavier oils from the refinery seldom
reach the relief headers.
7.3.4
Pipe Roughness (ε)
A default value of 0.0019 in (0.05 mm) is to be used in INPLANT for pipe roughness,
representing commercially shipped carbon steel pipe. Where a more appropriate
value is required, due to high corrosion rates, it is to be substituted.
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Hydraulic Evaluation
Evaluation of a flare system using INPLANT consists of six steps, as follows:
o
Gather flare system data such as UFDs, P&IDs, plot plans, isometric
drawings, and a PSV summary list.
o
o
Prepare a flare system sketch.
Use the flare system data to develop the required input data to the program.
o
Build and debug the model. Set up single link hydraulic systems to complete
analysis in sections where critical flow develops.
o
Review and document model results. Indicate flow and backpressure at
each PSV and other critical locations.
o
Identify any system deficiencies. Verify that the backpressure at the outlet of
each relieving PSV is acceptable.
An outline of this approach and a sample calculation is provided in Appendix B-8.
7.4
FLOW METERING
7.4.1
Design Discussion
Flare gas measurement is required for smokeless flaring control, monitoring
environmental impact or sometimes tracing leakage from process equipment.
There is always some flow through the relief headers which discharge to flares.
Ideally, only the purge gas passes through the system continuously. This is the
small flow of fuel gas or inert gas that is injected at various points in the system to
maintain a positive flow and to prevent air from entering. Actually, there is always
some leakage from imperfectly seating relief valves and vent valves.
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Higher flows, other than those during major upsets, do occur when relief valve
leakage is outside the acceptable limits or when a vent valve is left open
inadvertently. Such excessive flows can sometimes be detected by an increase in
the size or a change in color of the flame at the flare tip or by noise at the point of
leakage.
7.4.2
Methods
There are two special circumstances to take into consideration when devising a flare
meter. Firstly, the properties of the gas flowing in the header are generally not
known. Therefore, a chemical analysis, or at least a density determination, must
accompany the flow measurement in order to interpret it. Secondly, the linear
velocity in the flare header under normal operating conditions is extremely low,
requiring very sensitive detection techniques.
Two metering methods which may be considered are:
7.4.2.1
Flow Meter
A reasonably accurate measurement of the flare gas flow rate has been
difficult to achieve by conventional methods because of the very wide flow
or velocity range and constant change of the composition of the gas. In
the past annubar type instruments have been used but they have a limited
turndown ratio of about 10:1. The thermal type mass flow meters
manufactured by FCI does not provide the desired accuracy due to varying
gas composition. The thermal flow meters are more suitable for the
measurement of large flows in air, nitrogen or natural gas lines where
composition is relatively constant.
One of the major benefits of flare gas measurement is to determine the
amount of steam to be introduced into the flare, in order to achieve the
smokeless flaring. Also, it is important to minimize the lag time in the
introduction of the steam based on the flare gas flow measurement, rather
than wait until the smoke appears. Additionally, disproportionate amount
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of steam can be introduced in an untimely manner if the operator has to
rely only on the observation of the smoking flare.
Panametrics Model 7168 ultrasonic flow meter is specially designed to
measure the gas flow in headers up to 120 inches in diameter. It uses two
ultrasonic transducers located in the gas flow. Each of these transducers
located in the gas flow transmit and receive coded ultrasonic pulses.
Based on the measurement of time, it takes for an ultrasonic pulse to travel
from one transducer to the other, the instrument determines the velocity of
the gas in the flare header. The flow meter also measures the speed of
sound in the gas. The molecular weight of the flare gas is then determined
by the instrument using pressure, temperature and sonic velocity
measurements.
The basic features and specifications of the model 7168 flow meter are
summarized below:
•
Measures velocity, volumetric and mass flow rate
•
Measures instantaneous average molecular weight
•
No moving parts, no holes or tubes; tolerant to dirty or wet conditions
•
Mass flow rate range
0 to 4,409,200 lb/h (2,000,000 kg/hr)
•
Molecular weight range
0 to 60
•
Pressure
0 to 285 psig (19.6 barg, 20 kg/cm²G)
•
Temperature
- 58°F to 302°F (- 50°C to 150°C)
•
Velocity Range
0.1 to 280 ft/s (0.03 to 85 m/s)
•
Turndown
1000 : 1
•
Accuracy
5% or better
•
Input Power
110 or 220 VAC, 50/60 Hz or 12 or 24 V
DC at 25 watts
Ultrasonic transducers can be installed by use of hot or cold tap on an
existing flare system. The installation of the transducers is made through
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two 2” full opening valves. Similarly hot tapping is required for pressure
and temperature measurement instrumentation.
The vendor brochure data for Panametrics Model 7168 and FCI thermal
flow meters are attached. Currently Parametrics Model 7168 is the only
flow meter which is specially designed for the flare gas flow
measurements, and offers the accuracy for its usefulness in control of the
steam flow rate for smokeless flaring.
7.4.2.2
Dilution Method
In the dilution method a tracer gas is metered into the header. A sample is
taken downstream of the injection point and analyzed. The degree of
dilution of the tracer gas can then be used to compute the total flow.
Nitrogen is a suitable tracer gas. If nitrogen is already present in the flare
gas, the background nitrogen concentration is established and taken into
consideration in the computation. The sample point should be at least 20
to 30 pipe diameters downstream of the injection point to ensure proper
mixing. The injection rate must be high enough to result in at least 1 to 2
percent (by volume) nitrogen in the sample.
The advantage of this method is that it is inexpensive since the metering
device is reduced to a rotameter. One disadvantage is that for very low
flow rates considerable time is required for the tracer gas to travel from the
injection point to the sampling point. Since the flow rate in the header is
originally unknown, the operator does not know exactly when to take the
sample and he may have to take more than one.
If the flow is large, then the required rate of tracer gas injection can be
considerable. However, this rate does not have to be held very long
because a sample an be taken within minutes. The maximum tracer gas
injection capacity must be established in accordance with the top of the
range of flare flow to be measured. On the other hand, the total nitrogen
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supply which must be available for one measurement is set by the bottom
of the flow range to be covered.
7.5
SEALING AND PURGING
7.5.1. Sealing
7.5.1.1
Design Discussion
A flare system is subject to explosion hazards when air is present in the
system. Factors which contribute to air flow into the system are: thermal
contraction of hot gases in the system when flaring stops; oscillation
caused by rapid closure of a spring-operated safety relief valve; velocity of
air passing over the top of a stack.
When flaring stops, the hot gas cools rapidly and the contracting gas
creates a low pressure zone permitting the influx of air. Rainfall greatly
accelerates cooling.
Oscillation is caused by the rapid closure of a spring-operated safety relief
valve. The column of gas in the header continues to travel under its own
momentum and creates a low pressure zone behind it. As the column
recedes, air is drawn into the stack.
Wind over the top of a stack increases air flow with a resultant drop in
pressure, which causes gas to be drawn from the stack. Air flows in to
replace the gas at a rate that increases with higher wind velocities.
A water seal in a drum is to be provided near the stack or at its base and is
to be used in combination with a gas seal just below the stack tip. The gas
seal reduces the quantity of purge gas required.
Purge gas is required for all systems. In addition, a liquid seal, flame
arrestor, or gas seal must be provided.
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Gas Seals
Gas seals should be provided on all flues. Gas seal types can be divided
into two categories:
1.
Baffle type which is know under various trade names including
“Arrestor Seal” (John Zink Company), “Fluidic Seal” (National Airoil
Burner company) and “IG Series Seal” (Flaregas Corporation).
Figure 7.1 shows a baffle type of gas seal and its cutaway view. This
seal allows flow in one direction with little resistance. However,
reversal of flow results in very high resistance. This means that gas
can flow out of the flare stack but the counter flow of air back into the
stack is greatly restricted.
The baffle type of gas seal consists of a series of fixed baffles,
shaped like open-ended cones, mounted within a flare tip. Each
succeeding baffle of the seal encountered by purge gas traveling up
the stack has a larger aperture than that below it. Air attempting to
enter the stack is turned back against itself by the first baffle. Each
succeeding baffle, with its progressively smaller aperture, further
reduces the flow of air.
2.
Labyrinth type which is know under various trade names including
“FX Series Seal” (Flaregas Corporation) and “Molecular Seal” (John
Zink Company).
Figure 7.2 shows a labyrinth type of seal and its cut-away view. This
seal is installed below the flare tip and connected to both the stack
and the tip through flanged joints. The labyrinth seal hinders the flow
of air into the stack whether the stack gas is lighter or heavier than
air. Assuming no purge gas is used and the gas in the stack is lighter
than air, the atmospheric air enters zone A due to gravity. However,
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the air is trapped in Zone B because it is heavier than the gas in this
region. Atmospheric air can penetrate Zone B only through
molecular diffusion. If the gas in the stack is heavier than air, the
atmospheric air cannot penetrate Zone A due to its gravity, except by
molecular diffusion. Because it functions on the basis of molecular
weights different from air (heavier and lighter), this type seal is
frequently referred to as a “Mole Seal”.
The water drain from the labyrinth seal must be looped at grade to
provide a 10 feet (3 meters) seal or two times the maximum operating
pressure at the base of the stack, whichever is greater. The drain is
piped to the knockout drum. A water supply must be available to
maintain the water seal level.
In general, the baffle seal costs less than the labyrinth seal because
of its simple construction and light weight. Less structural support is
required for the baffle seal than for the labyrinth seal.
The procedure for estimating pressure drop across either type of seal
is the same as that for a flare tip. Refer to Section 9.0 for a
description of the method.
7.5.1.3
Water Seals
o
Nomenclature
Unless otherwise stated, all symbols used in this section are defined
as follows:
CHAP7-R1.DOC
DD
=
Inside diameter of water seal drum, ft (m)
DP
=
Inside diameter of main vapor inlet line to drum, ft (m)
HL
=
Depth of water below end of main vapor inlet line, ft (m)
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HS
=
Depth of water above end of main vapor inlet line, ft (m)
HV
=
Height of vapor space above the water seal, ft (m)
Design Procedure
Water seal drums serve as backup to the purge gas system.
However, they may also be used to divert flow during normal plant
operation to a vapor recovery system or small relief flows to a ground
flare. Any larger streams would discharge to an elevated flare with
purge gas being introduced downstream of the water seal drum.
The water seal drum is often a vertical vessel installed at the base of
a flare stack. In this instance, the drum also provides structural
support for the flare stack. The water seal may also be provided in a
vertical or horizontal drum separate from the flare stack or ground
flare.
Figure 7.3 shows a vertical water seal drum and its design features.
They also apply to the design of a horizontal water seal drum except
that the main vapor inlet connection is located on the top of the drum.
Design features include:
a)
The water seal depth is normally maintained at 12 in (30 cm)
above the bottom of the inlet line to keep air out of the system
b)
Perforated, anti-sloshing baffles inside the drum prevent
pulsating combustion in the flare and subsequent sound
pulsation which may be annoying to people.
c)
CHAP7-R1.DOC
Sufficient vapor space is provided above the water seal to
allow for the separation of the entrained seal liquid.
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The volume of water above the end of the inlet line must be
sufficient to fill 10 feet (3 meters) of the line. This prevents
flashback into the line.
e)
The height of the water seal leg from the drum bottom is
based on preventing vapor release from the drum bottom
during maximum flaring conditions.
The water level in the drum is maintained by a constant supply of
water. The relief gas inlet pipe is projected into the drum and
immersed in the water to form a positive seal. When the facility is
located in the cold climate, insulation and steam tracing is used to
avoid freezing of water. Water from the drum is normally discharged
to the oily water sewer system. However, drum effluent water may
require some other type of treating, such as, hydrogen sulfide
removal, prior to discharge from the plant or reuse within the plant.
The drum may require provisions for skimming hydrocarbon liquids
from the surface of the water. Removal can be either on a
continuous or an intermittent basis.
If the drum is located at the base of the flare stack, the drum size
may have to be increased to meet structural support requirements for
the stack. These are determined by the structural engineer.
See Chapter 8.0 for details for engineering design of the water seal
drum.
7.5.2
Purge Gas
Unless otherwise stated, all symbols used in this section are defined as follows:
V =
Total volume of main flare header and knockout drum, ft3 (m3 )
Ta =
Ambient temperature, °R (°K)
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FLARE SYSTEM
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T
=
Flare gas temperature in main header, °R (°K)
t
=
Cooldown time, h
Q =
7.5.2.1
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Purge gas flow, SCFH (standard m3/h)
Purge Gas Control System
The flare system must be continuously purged with fuel gas or inert gas to
prevent explosion hazards due to air entering the system. The purge gas
is introduced at the start of each branch header. Each purge gas
connection to a branch header includes a flow indicator and ball valve. A
check valve is also required when inert gas is used as the purge gas. The
minimum purge gas line size is one inch.
A purge gas connection is recommended on the main header to
supplement the continuously introduced purge gas. This connection is
located downstream of the tie-in between the main header and the last
branch header handling relief gases hotter than 150 °F (66 °C). The
additional purge gas is required immediately after gases with a
temperature higher than 150 °F (66 °C) have been discharged to the flare
in order to prevent air entering the system due to thermal contraction. This
special connection includes a flow control valve, flow indicator, and block
valve. A check valve is also required when inert gas is used as the purge
gas.
The supplemental purge gas flow is controlled by the gas temperature in
the main header. The control valve is automatically closed after
supplemental purge gas has been injected for a minimum of 15 minutes. A
local indicating light and a remote alarm are provided to inform the
operators that the valve is open. The items furnished by a flare vendor
include a control panel, timer, high temperature sensor, low pressure
sensor, and the local indicating light. The instrumentation provided by the
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engineering contractor includes the remote alarm, flow control valve, and
flow indicator.
A typical purge gas supply piping configuration is shown in Figure 7.4
7.5.2.2
Normal Purge Gas Requirements
When the purge gas is methane or a heavier gas and when the wind
velocity is 13 ft/s (4 m/s) or less, the total normal purge gas flow required
from the branch header connections is as follows:
a)
For an elevated flare: 0.1 ft/s (0.03 m/s) velocity in the flare stack or
tip, whichever has the greater diameter. The quantity of purge gas
required depends on the type of seal selected. This velocity is 3
times higher than the stated requirement for a labyrinth type seal
from vendors.
b)
For ground flares: 0.1 ft/s (0.03 m/s) velocity in the first stage header.
Flow requirements for purge gases lighter than methane are three times
those for methane. For wind velocities above 13 ft/s (4 m/s), purge gas
requirements increase as the square root of the wind velocity.
7.5.2.3
Upset Condition Purge Gas Requirement
The upset condition purge gas flow required from the special main header
connection in addition to the normal purge from the branch header
connections is estimated as follows:
CHAP7-R1.DOC
a)
For an elevated flare regardless of the type of seal used: 3 ft/s (1
m/s) velocity in the flare stack or tip, whichever has the greater
diameter.
b)
For a ground flare: 3 ft/s (1 m/s) velocity in the main header.
FLARE SYSTEM
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In the final design, the flow can be calculated more precisely by the
formula below:
Q = [V/t][ 1 - (Ta/T)]
7.5
Use 0.25 hour (15 minutes) for cooldown time unless otherwise indicated
by calculations. The maximum flow is determined using the minimum
ambient temperature and the maximum flare gas temperature.
7.5.2.4
Startup/Shutdown Purging
Start-up and shutdown of the entire flare system cannot be performed
without shutting down and draining all process equipment and piping. This
is not normally done at any time in the operating life of the plant. Units,
however, are shutdown on a regular basis for various reasons such as
annual safety inspections, replacement of catalyst charges, etc. In order to
enable the initial startup and subsequent shutdowns and restarts to be
performed, vents are usually provided at various locations in the piping.
See Figure 7.4 for details. Note the unit header block valves with blinds.
o
Start-up
Use nitrogen or steam to rid the relief headers of air. Vent the vapors
to atmosphere until the test for oxygen in the piping is acceptable.
The headers can be swept clear with continuous flow or pressured up
to about 50 psig (3.4 barg, 3.5 kg/cm2 G) and depressured to near
atmospheric pressure in a cycle. The cost of inert media should be
used as the basis to select between sweeping and pressure cycling.
o
Shutdown
Use nitrogen to clear the header or subheader by pushing the
hydrocarbons towards the flare. Like the startup case, this can be
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done continuously or by pressuring and depressuring the subheader.
Continue purging until the gas is well below the lower explosion limit
(LEL). Block in the unit header from the flare main and close the
blind. Purge the unit header with air if personnel entry is required.
The header or subheader may require steam out if any tars or heavy
oils are present. This requirement should be based on prior
experience with the service.
7.6
REFERENCES
1.
Reed, R.D., “Furnace Operations”, Gulf Publishing, 12-31 (1973).
2.
Reed, R.D., “What is Flare’s Proper Purge Rate?”, Oil and Gas Journal (February 14,
1972).
3.
Seebold, J. G., “Reduce Noise from Pulsating Combustion in Elevated Flares”,
Hydrocarbon
Processing, 225-227 (September, 1975).
4.
Trade literature from the following flare design specialists:
•
•
CHAP7-R1.DOC
John Zink Company
Kaldair
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TABLE 7.1
PSV INLET/OUTLET CALCULATIONS, DESIGN CRITERIA
A.
Inlet Calculations
Valve Type
Phase
Max Loss (1)
(% Set P)
Conventional
Conventional
Conventional
VAP
LIQ
VAP/LIQ
3
3
3
Rated @ 10% overpressure
Required flow (6)
Rated @ 10% overpressure
Balanced Bel
Balanced Bel
Balanced Bel
VAP
LIQ
VAP/LIQ
3
3
3
Rated @ 10% overpressure (2)
Required flow (6)
Rated @ 10% overpressure (2)
Nonmod Pilot
Nonmod Pilot
Nonmod Pilot
VAP
LIQ
VAP/LIQ
3
3
3
Rated @ 10% overpressure (2) (4)
Required flow (4) (6)
Rated @ 10% overpressure (2) (4)
Mod Pilot
Mod Pilot
Mod Pilot
VAP
LIQ
VAP/LIQ
3
3
3
Required flow (4)
Required flow (4) (6)
Required flow (4)
B.
Flow Rate Basis
Outlet Calculations (5)
Valve Type
Phase
Max Loss (1)
(% Set P)
Conventional
Conventional
Conventional
VAP
LIQ
VAP/LIQ
10
10
10
Balanced Bel
Balanced Bel
Balanced Bel
VAP
LIQ
VAP/LIQ
50 (3)
50 (3)
50 (3)
Required flow (7)
Required flow (6) (7)
Required flow (7)
Nonmod Pilot
Nonmod Pilot
Nonmod Pilot
VAP
LIQ
VAP/LIQ
50 (3)
50 (3)
50 (3)
Required flow (7)
Required flow (6) (7)
Required flow (7)
Mod Pilot
Mod Pilot
Mod Pilot
VAP
LIQ
VAP/IQ
50 (3)
50 (3)
50 (3)
Required flow (7)
Required flow (6) (7)
Required flow (7)
CHAP7-R1.DOC
Flow Rate Basis
Rated @ 10% overpressure (8)
Required flow (6)
Rated @ 10% overpressure (8)
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TABLE 7.1 (continued)
PSV INLET/OUTLET CALCULATIONS, DESIGN CRITERIA
NOTES:
1. Inlet loss criteria is independent of actual allowable overpressure. However, if overpressure
exceeds 10%, valve capacity should be adjusted for pressure loss in excess of 3%.
2. Rated flow for inlet calculations should be based on no built-up back pressure, since valve
chattering could occur before the header is pressurized.
3. Valve capacity may be reduced when backpressure exceeds 30% of set pressure. Reduction of
capacity (Kb factor) is shown in Figure 5.14. Valve manufacturer should always be consulted
when back pressure exceeds 50% of set pressure.
4. For pilot valves with a remote tap, inlet drop is measured only from the protected equipment to the
tap location. For flowing pilot valves, the pilot line should be large enough to make the pressure
drop in the pilot line negligible. If a pilot tap is placed in a flowing line, the tap should be a pitot
type connection so that the velocity head is recovered. Valve capacity should be adjusted for
pressure drop to the main valve inlet in excess of 3% of set pressure.
5. Maximum back pressure based on valves sized for 10% overpressure. For valves designed for
greater % overpressure, maximum back pressure must also be evaluated for the rated flow at the
higher % overpressure conditions. At this condition, maximum % back pressure must not exceed
% overpressure.
6. Incompressible flow is normally self-limiting based on pump and/or system hydraulics. Pressure
will drop off if the valve opens further than needed to pass the required relieving rate.
7. Although actual valve flow may exceed required flow, added backpressure will act to limit flow and
the valve performance will be stable.
8. Use rated flow only for individual relief valve outlet lateral. Once piping diameter has increased to
become a combined lateral or subheader, hydraulics should be based on required flow.
9. Use rated flow only fr individual relief valve inlet piping. When a common header is installed for
multiple valves, use required flow for the common header hydraulics.
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TABLE 7.2
ORIFICE/INLET AREA RATIO FOR
STANDARD RELIEF VALVES
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TABLE 7.3
MAXIMUM ALLOWABLE EQUIVALENT LENGTHS
OF INLET PIPING TO COMPLY WITH 3%
INLET LOSS CRITERIA FOR RELIEF VALVE(1)
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TABLE 7.4
ORIFICE/OUTLET AREA RATIO
FOR STANDARD RELIEF VALVE
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FLARE SYSTEM
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TABLE 7.5
TYPICAL OUTLET NOZZLE LENGTHS
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TABLE 7.6
TYPICAL REDUCER ANGLES
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TABLE 7.7
TYPICAL WELD NECK FLANGE LENGTHS
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TABLE 7.8
TYPICAL FRICTION (f ) FACTORS
(Moody or Darcy-Weisbach)
FOR CLEAN CARBON STEEL PIPE
(εε = 0.05 mm)
NOMINAL PIPE SIZE
INSIDE DIAMETER
FRICTION FACTOR
(NPS)
(inches)
f
2" SCH 40
3" SCH 40
4" SCH 40
6" SCH 40
8" SCH 40
10" SCH 40
12" SCH 40
14" ST WALL
2.067
3.068
4.026
6.065
7.981
10.020
12.000
13.250
0.0190
0.0175
0.0160
0.0150
0.0140
0.0135
0.0129
0.0126
16" ST WALL
18" ST WALL
15.250
17.250
0.0123
0.0120
20" ST WALL
24" ST WALL
30" ST WALL
19.250
23.250
29.250
0.0119
0.0115
0.0110
36" ST WALL
35.250
0.0107
Notes: 1. The friction factors above apply at high Reynolds numbers (fully developed
turbulent flow); above 2 x 105 for 2" increasing up to 1 x 106 for 24" and larger.
Values above these Reynolds numbers are typical for flare headers and laterals.
Source: Crane Co., Technical Paper No. 410M, 1986, pg. A-23
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FIGURE 7.1
BAFFLE TYPE SEAL
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FIGURE 7.2
LABYRINTH TYPE SEAL
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FIGURE 7.3
VERTICAL WATER SEAL DRUM
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FIGURE 7.4
FLARE PURGE GAS SUPPLY
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FIGURE 7.5
TYPICAL PRESSURE RELIEF VALVE INSTALLATION:
ATMOSPHERIC (OPEN) DISCHARGE
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FIGURE 7.6
TYPICAL PRESSURE RELIEF VALVE INSTALLATION:
CLOSED SYSTEM DISCHARGE
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FIGURE 7.7
TYPICAL PRESSURE RELIEF VALVE
MOUNTED ON PROCESS LINE
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FIGURE 7.8
TYPICAL PRESSURE RELIEF VALVE
MOUNTED ON LONG INLET PIPE
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FIGURE 7.9
TYPICAL PILOT-OPERATED PRESSURE RELIEF
VALVE INSTALLATION
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FIGURE 7.10
TYPICAL RUPTURE DISK ASSEMBLY INSTALLED
IN COMBINATION WITH A PRESSURE RELIEF VALVE
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FIGURE 7.11
TYPICAL PRESSURE RELIEF VALVE INSTALLATION
WITH AN ISOLATION VALVE
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KNOCKOUT, BLOWDOWN, SEAL,
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SECTION 8.0
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KNOCKOUT, BLOWDOWN, SEAL, QUENCH DRUMS AND PUMPS
Four types of vessels or drums are typically used in the design of flare systems for specific
reasons:
o
Knockout drum (Liquid/Gas Separator)
o
Blowdown drum (Liquid holdup)
o
Seal drum (To prevent the air ingression into the flare system to avoid explosion)
o
Quench drum (Condenser/Cooler)
The types of drums required to suit the purposes are outlined below.
All four types of drums are to be designed for a minimum pressure of 50 psig (3.4 barg, 3.5
kg/cm2 G) to withstand internal pressure rises associated with the unforeseen combustion
of hydrocarbons with the ingressed air. Design temperatures for each drum should be
consistent with maximum and minimum temperatures for material which can flow to the
drum, including consideration of boil-off temperature for accumulated material. Heat
transfer calculations can sometimes be used to moderate these temperatures based on
heat loss or gain from the point of relief through the piping upstream of the drum.
8.1
KNOCKOUT DRUM
8.1.1
Purpose
A knockout drum is required where sufficient amounts of hydrocarbon liquid are
entrained with or condensed from the gas to avoid possible fire hazards from liquid
droplets falling out of the flare. Industry and vendor data indicate that the removal
of liquids having a particle diameter of 400-600 microns is required to prevent such
occurrences. Therefore, the recommended knockout drum sizing formulas are
based upon slowing down the vapors enough to allow gravity settling of liquid
droplets of that size. In the size range of 600-1000 microns a significant increase in
smoking can result. Beyond 1000 microns, carryover of flaming particles can occur.
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For new drums, the design particle size should be 400 microns. For revamp work, a
maximum particle size of 600 microns is recommended. However, carryover of
particles slightly in excess of 600 microns in an existing K.O. drum may be
determined to be acceptable on careful review if the impact and/or probability of
occurrence is low.
The function of the knockout drum(s) may be combined with that of the blowdown
drum to provide a single drum, where small quantities of liquid are involved.
Consider the impact of this selection on the slope of the flare header(s) and project
economics.
The knockout drum is normally located close to the flare stack. For this reason,
consider the amounts of liquid that may be required to flow though the gas flare
header if local or area blowdown drums are not used. The liquid material collected
in the knockout drum is usually pumped back to the slop tanks.
For segregated cold and warm flare headers, separate dry and wet flare knockout
drums are provided due to different piping material’s requirement.
Figure 8.1 provides a typical sketch of a horizontal knockout drum.
8.1.2
Design Parameters
The knockout drums may be either horizontal or vertical. Horizontal drums are more
common for large relief loads for the following reasons:
o
The required elevation of the relief header is lower than for a vertical drum.
o
The horizontal drum would cost less than the vertical drum which has
equivalent capacity.
Unless otherwise stated, all symbols used in this section for sizing horizontal/vertical
drums are defined as follows:
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AT
=
Total cross section of the drum, ft2 (m2)
AV
=
Cross sectional area for vapor flow, ft2 (m2)
AL
=
Cross sectional area for liquid inventory, ft2 (m2)
2
C(Re) =
Characterization parameter for drag coefficient
C
=
Drag coefficient from API RP 521, 1990, section 5.4.
See Figure 8.2
Ud
=
Droplet settling velocity, ft/s (m/s)
Uv
=
Average vapor velocity, ft/s (m/s)
µ
=
Vapor viscosity, centipoise
MW
=
Molecular Weight of the vapor
T
=
Temperature of the relieving vapors, °F (°C)
Tabs
=
Absolute temperature, °R (°K)
f
=
Mass fraction of total vapor flow rate condensed
g
=
Acceleration due to gravity, 32 ft/s/s (9.8 m/s/s)
Dp
=
Droplet diameter, m, typically 400 to 600 µ (10-6 m)
Dd
=
Drum inside diameter, ft (m)
Rd
=
Drum inside radius, ft (m)
Lqs
=
Liquid storage capacity, ft3 (m3 )
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Liquid storage capacity equivalent to 15 minutes at the maximum
instantaneous liquid carryover rate is required when the knockout drum
pump or pumps start automatically.
Liquid storage capacity equivalent to 30 minutes at the maximum
instantaneous liquid carryover rate is required when the knockout drum
pump or pumps start manually.
When there is no clearly defined liquid carryover rate, the maximum
operating liquid level in the drum is assumed to be 25 percent of the
drum diameter for a horizontal drum or 20 percent of the tangent to
tangent height for a vertical drum.
ρl
=
Density of the liquid at operating conditions, lb/ft3 (kg/m3)
ρv
=
Density of the vapor at operating conditions, lb/ft (kg/m )
Qv
=
Vapor flow rate, ft /s (m /s)
Ql
=
Liquid flow rate, ft3/s (m3/s)
θ
=
Liquid drop out time, s
hv
=
Vertical drop available for liquid drop out, ft (m)
hL
=
Liquid depth, ft (m)
L
=
Assumed Length of the horizontal drum for disengagement, ft (m)
Lreq
=
Required Length of the horizontal drum for disengagement, ft (m)
W
=
Total mass flow rate, lb/h (kg/h)
Wv
=
Vapor flow rate, lb/h (kg/h)
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WL
=
Liquid flow rate, lb/h (kg/h)
P
=
Operating pressure, psig (barg, kg/cm2 G)
Pabs
=
Absolute pressure, psia (bara, kg/cm2 A)
R
=
3
Ideal gas constant, 10.732 psi-ft /lbmol-°R
(0.08315 bar-m3/kmol-K, 0.08479 (kg/cm2)-m3/kmol-K)
t
=
time, h (See notes associated with symbol Lqs)
Va
=
Allowable vapor velocity, ft/s (m/s) (For vertical
knockout drum)
8.1.2.1 Horizontal Drum
See Figure 8.1 for a typical horizontal knockout drum.
The following procedure (for API RP-521, Third Edition, Paragraph 5.4.2.1)
should be followed to determine the size of a horizontal drum. A
spreadsheet program to perform this calculation is provided in Appendix D
and a sample calculation using that program is provided in Appendix B-5.
Calculate the vapor density if not known.
ρv = ( Pabs MW ) / ( R Tabs)
8.1
2
8
2
3
C(Re) = 0.95 x 10 ρv Dp (ρl - ρv )/µ
2
8
C(Re) = 0.13 x 10
ρv Dp3
(ρl - ρv )/µ
2
8.2 (English)
8.2 (Metric)
Calculate C(Re)2 using Equation 8.2 and obtain C by interpolation of Figure
8.2.
Ud = 1.15 [(g Dp (ρl - ρv ))/( ρv C )]0.5
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Qv = W(1-f) / (ρv x 3600)
8.4
Ql = Wf / (ρ l x 3600)
8.5
W L = Wf
8.6
Lqs = W L t / ρ l
8.7
Assume drum diameter Dd and try the length of the drum, L = 2.5 Dd or 3 Dd
AL = Lqs/L
8.8
AT = πRd2
8.9
Calculate AL/AT and obtain hL/Dd from Table 8.1.
hv = Dd - hL
θ
= hv / Ud
8.10
Av = AT - AL
8.11
Uv = Qv / Av
8.12
Lreq = Uv θ
8.13
(Note: Split flow does not change value of Lreq, since both vapor velocity
and travel distance are cut by half).
If Lreq is greater than L, increase the drum diameter and repeat the
calculations again until Lreq is less than L. (Note: If assumed drum diameter
equals maximum shipping diameter, then L must be increased and the value
of L/D used following Equation 8.7 may need to be greater than 3).
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Vertical Drum
The following procedure should be followed to determine the size of a
vertical knockout drum.
Calculate ρv from Equation 8.1, if unknown.
Calculate Va = Ud (from Equation 8.3)
Calculate Qv = W (1 - f) / (3600ρv ) (from Equation 8.4)
Dd = 1.128 ( Qv /Va )0.5
8.14
Provide a minimum vapor space above the liquid level equal to one meter or
one drum diameter, whichever is greater.
8.1.3
Design Details
Knockout drum(s) are to be provided with a liquid level indicator, level alarm and an
automatic liquid drawoff, as required by client’s design guidelines and/or standards.
A steam coil (or heater) may be required to allow for winterization and weathering of
the liquid from drains and emergency liquid releases prior to disposal. A small
quantity of steam should be allowed to pass through the coil continuously for
winterization and it could be increased for weathering when required.
An internal baffle or a 90° elbow is often used at the inlet to a relatively small
horizontal/ vertical drum to direct the flow either downwards or sideways to knockout
larger liquid droplets. However, the effectiveness of this arrangement is not
theoretically proven for horizontal drums. Installation of internal baffles for a
horizontal K.O. drum is not recommended, unless otherwise requested by the client.
An internal baffle is still recommended for a vertical K.O. drum as described in the
Process Design Manual 225-006, Section 4.4.
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A water boot should be provided at the bottom of a horizontal drum to remove
separated water from the liquid hydrocarbons, if considerable water accumulation is
anticipated.
The INPLANT program (from Simulation Science, Inc.) can be used to calculate
header heat losses and resultant condensation rates, if required, to establish a
rational basis for liquid hydrocarbon and water collection rate in the knockout drum.
Detailed stream composition input must be required to complete this calculation.
8.2
BLOWDOWN DRUM
8.2.1
Purpose
The purpose of a blowdown drum is to capture sizable liquid releases from the
process. These releases may be intentional, such as equipment drainage through
the refinery liquid drain system during shutdown operations or of an emergency
nature, such as relief or safety relief valve discharges. The drum must be capable
of accepting the anticipated liquid loads without filling beyond the maximum
operating level of the drum. This maximum operating level must still provide a gas
flow path with sufficiently low gas velocity to allow gravity separation and prevent reentrainment beyond the design droplet size. Blowdown drums tend to be located
near the sources, close to the battery limits of a process unit or area to reduce the
amount of piping subject to two phase slug flow. It is not a good practice to design
for significant amounts of liquid to run long distances through the plant in the vapor
flare header.
Figure 8.3 provides a sketch of a typical horizontal blowdown drum.
Where the liquid flow quantities anticipated are relatively low, a knockout drum will
be sufficient and a blowdown drum may not be required.
Blowdown drums may be advantageous, even for relatively low liquid flow quantities
for the following services:
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To collect highly viscous liquids in a relatively small vessel so that the relief
header will not become fouled or plugged and the main knockout drum need
not be designed to handle this liquid.
8.2.2
o
To collect liquids which should not be mixed with other streams due to
economic consideration or corrosion potential.
o
To collect liquids within process areas where continuous sloping of the
headers toward the main knockout drum is not feasible.
Design Parameters
Blowdown drums may be either horizontal or vertical. Horizontal drums are more
common since the diameter of vertical drums becomes prohibitive in the case of
large flare relief flows. Also, the higher elevation of vertical drums increases the
height of the flare headers required to provide free drainage of the header . The
higher elevation of the flare headers increases the capital investment due to the
higher cost of piping supports and structures.
The horizontal and vertical blowdown drums are designed in the same manner as
knockout drums (see Section 8.1.1.2). The primary difference is the magnitude of
the design liquid volume.
8.2.3
Design Details
Blowdown drum(s) are to be provided with a liquid level indicator, level alarm and an
automatic liquid drawoff as required by client’s design guidelines.
A steam coil may be required to allow for winterization and weathering of the liquid
from drains and emergency liquid releases prior to disposal.
Where several process units or areas use a common blowdown drum, a combined
dedicated liquid relief header from these areas to the blowdown drum should be
provided. The separated vapor can then be combined with the vapor relief header.
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The refinery liquid drain system(s) usually flows to this drum. Provide a 4” or 6”
connection on the drum for this service early in the project development. The final
size can be adjusted as the project develops.
Allowance must be made for retention of the following amount of liquid material:
o
Liquid drains per the project standard
o
Emergency releases - maximum volume relieved over 30 minute period
The horizontal blowdown drum should be sloped 0.33 in per 3.3 ft (0.83 cm per
meter) towards the liquid outlet nozzle. A water boot may be required at the bottom
of the horizontal drum to remove separated water from the liquid hydrocarbons. The
liquid material is usually disposed of to the slop tanks following weathering or
directly to refinery process unit. The inlet or outlet nozzles may be installed at an
angle from the vertical if the elevation of the flare header is otherwise too low with
respect to the blowdown drum inlet nozzles.
Internal baffles are required at the inlet to a horizontal/vertical drum to direct the flow
either downwards or sideways to knock out larger liquid droplets.
The blowdown drum should have a relieving device if it can be subjected to
overpressure during an external fire case.
8.3
SEAL DRUM
8.3.1
Purpose
A seal drum provides a positive seal against the ingression of atmospheric air from
the flare stack into the flare header. The reason for this seal is to prevent, as much
as possible, the development of a combustible air/hydrocarbon mixture and thus
prevent flashback in the flare system.
See Section 7.5.1 for a detailed discussion of the seal drum. Also, refer to Figure
7.3, Vertical Water Seal Drum and Figure 8.4, Horizontal Seal Drum. It is a
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common practice to install an integral vertical seal drum at the base of a self
supported flare stack.
The decision to select a horizontal or vertical seal drum should be made on the
basis of the total installed cost. Large drum diameters will lead to the need for field
fabrication.
Another essential purpose for seal drums is to stage the flows to several
destinations such as:
o
A ground flare and an elevated flare (see Figure 8.5)
o
An incinerator and an elevated flare.
o
An operating smokeless flare and an emergency non-smokeless flare (see
Figure 8.6)
The second destination is sealed from the flare header until the header
backpressure is sufficient to overcome the water seal head.
8.3.2
Design Parameters
8.3.2.1
Vertical Water Seal Drum
For a vertical water seal drum, the drum diameter (Dd) and the tangent to
tangent height are set by the diameter of the relief header (Dp) entering the
seal drum. Carryover of sizable droplets of the seal water is not a problem
with the flare. If a minimum backpressure is required upstream of the seal
drum (for flare gas recovery or flare staging), this backpressure will set the
height, h, of the inlet pipe to be submerged and can be calculated as
follows:
CHAP8-R1.DOC
h = 144 P / ρw
8. 15 (English)
h = 10,200 P / ρw
8. 15 (Metric)
h = 10,000 P / ρw
8. 15 (Metric)
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Where:
h
=
water seal depth, ft (m)
P
=
Required backpressure, psig (barg, kg/cm2 G)
ρw
=
Water density, lb/ft (kg/m )
3
3
Although the sizing basis should be verified and confirmed by the selected
flare vendor, the following general guidelines are useful in establishing the
drum diameter and height as given in Figure 7.3:
•
a)
2 x Dpipe
b)
1.414 x Dpipe + 4.0 ft (1.22 m)
c)
Diameter required to provide sufficient water to fill 10 ft (3 m)
of the inlet line with 1 ft (0.3 m) seal depth.
Vapor space height (Hv) = greater of:
•
•
•
CHAP8-R1.DOC
Drum diameter (Dd) = greatest of:
a)
1/2 Dd
b)
2 Dpipe + 3 ft (1 m)
Dip leg clearance (Hi) = greater of
a)
1 ft (0.3 m)
b)
0.25 Dpipe
Seal height = greater of
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a)
10 ft (3 m) water pressure gauge
b)
200% of maximum drum operating pressure
Where:
8.3.2.2
SECTION 8.0
Dpipe = Inlet pipe diameter, ft (m)
Horizontal Water Seal Drum
The design parameters and the general guidelines associated with a
horizontal drum are given in API RP 521 (Third Edition), paragraph 5.4.2.2.
Basic requirements are similar to those for vertical drums provided in
Section 8.3.2.1.
8.3.3
Design Details
The seal water level shall be maintained at the proper level automatically by liquid
makeup and proper drawoff provisions to assure protection against flashback. A
steam coil may be required for winterizing in the areas where the minimum
atmospheric temperature is lower than 32 °F (0 °C). Provide for continuous or
intermittent oil skimming to prevent hydrocarbon accumulation.
All water seal drums shall be supplied with a drain, pressure indicator (PI) at gas
inlet, temperature indicator (TI), utility connection for the fill line and a 20” minimum
manway.
Anti-slosh baffling should be added as required. See Figure 8.4 for horizontal
drums and Figure 7.3 for vertical drums.
8.4
QUENCH DRUMS
8.4.1
Purpose
A quench drum can be used to condense the vapor discharge from a relief device
for either later return back into the process after the relieving condition has passed
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or for disposal to the sewer. Generally, a quench drum is provided whenever the
material being relieved is too valuable to be burned in a flare or too toxic to be
relieved to atmosphere in a vent stack. Also, a quench drum can be used to cool
hot material so that the entire relief system does not need to be designed for the
higher temperature. A third purpose for a quench drum is to quench runaway
reactions. This last purpose is beyond the scope of this manual. Design for
runaway reaction quench is a specialized engineering task that must be treated
separately, usually by third party licensers. Refer to publications of the Design
Institute for Emergency Relief Systems (DIERS) of the American Institute of
Chemical Engineering (AIChE) for further information.
See Figure 8.7, Typical Quench Drum (Condensables) and Figure 8.8, Typical
Quench Drum (Emergencies).
8.4.2
Design Parameters
A quench drum is a direct contact condenser utilizing water or another suitable
quenching fluid flowing down across either baffle trays or disk-and-donut trays to
condense the rising vapors. A temperature sensing element on the vapor inlet
signals a quick opening valve on the cooling water or fire water line to flood the
tower with water. Sizing procedures for either a disk-and-donut tray column or a
baffle tray column are given in the following articles:
1.
“How to Size Shower Deck Baffled Towers Quicker”
o
Part 1, “Tower Diameter”, by A.D. Scheiman, March, 1965,
Petro/Chem Engineer, Page 28
o
Part 2, “Tower Tangent Length”, by A.D. Scheiman, April, 1965,
Petro/Chem Engineer, Page 75.
2.
CHAP8-R1.DOC
“Designing Direct Contact Coolers/Condensers” by J.R. Fair, Chemical
Engineering, June 12, 1972, page 91.
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Both procedures given above should be explored to establish the best design
economics.
8.4.3
Design Details
The quenching fluid in a quench drum or tower may be water, gas oil, or another
suitable liquid. The quench fluid and absorbed gas liquor collects in the bottom of
the drum for later removal. Quench requirements may be monitored by a
temperature or flow switch in the relief discharge header. The level of liquid in the
bottom of the drum or tower may be automatically controlled. Recovered liquid may
be cooled and recycled, dumped to the sewers, or sent to equipment for recovery of
the gas liquor. Alarms may be provided to signal operators in the event that design
liquid levels are exceeded. Steam coils must be provided in the liquid collection
zone to winterize the drum.
8.5
PUMPS
In most instances, pumps are provided to remove the liquids from blowdown, knockout and
quench drums. The pumps normally discharge to slops tanks. Occasionally, the liquid may
be recycled to the process or disposed of in a burn pit or incinerator.
The minimum required pump flow capacity is based on the largest of the following criteria:
o
3
3
425 ft /h (12 m /h)
o
150 percent of the maximum continuous (planned for longer than 30 minutes) liquid
relief. Typically, this is based on the expected condensation rate during a
continuous flaring contingency.
o
25 percent of the maximum instantaneous liquid relief load, i.e. pumpout of liquid
from a 15 minute relief event within one hour.
A spare pump is normally provided for this service which can be used in case the liquid
level rises too fast in the drum due to unforeseen circumstances or if the operating pump is
out of service. It is preferable to instrument the pump for automatic start and automatic
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stop with provision for manual override. Insurance industry representatives have in the past
requested this arrangement. However, careful evaluation should be conducted prior to
operating in the autostart mode to assure that any material which could accumulate in and
be pumped out of the drum can be accommodated at the destination (slop tank or
alternative location). Problems have occurred in some systems when material containing
light ends is pumped into a hot liquid and flashes, resulting in overpressuring of the
receiving tank. Hot and cold fluids can also pose potential handling concerns.
Horizontal centrifugal pumps are commonly selected for flare knockout drum service.
Centrifugal pumps are usually less costly and more reliable as well as being simpler to
maintain than positive displacement pumps. A pump engineer and/or pump vendor should
be contacted to review the selection.
The motor driven blowdown and knockout drum pumps should be connected to the
emergency power system or powered from two separate feeders. Where this is not
possible, a steam turbine driver should be provided to the spare pump to assure its
operation during partial or total power failure.
8.6
REFERENCES:
1.
Perry, R.H., and Chilton, C.H., Chemical Engineers Handbook (pp. 5-61 through 565), Fifth Edition, 1973, McGraw-Hill, New York.
2.
Lapple and Shepard, Industrial and Engineering Chemistry, 32, pg. 605, 1940.
3.
Zeng and Othmer, “Fluidization and Fluid-Particle Systems”, pp. 216-220, Reinhold,
New York, 1960.
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TABLE 8.1
TABLE OF GEOMETRY FOR CIRCLES AND ARCS
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TABLE 8.1
TABLE OF GEOMETRY FOR CIRCLES AND ARCS
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TABLE 8.1
TABLE OF GEOMETRY FOR CIRCLES AND ARCS
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TABLE 8.1
TABLE OF GEOMETRY FOR CIRCLES AND ARCS
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FIGURE 8.1
TYPICAL HORIZONTAL KNOCKOUT DRUM
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FIGURE 8.2
DRAG COEFFICIENT, C
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FIGURE 8.3
TYPICAL HORIZONTAL BLOWDOWN DRUM
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FIGURE 8.4
TYPICAL HORIZONTAL SEAL DRUM
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FIGURE 8.5
SCHEMATIC FOR COMBINED GROUND
FLARE AND ELEVATED FLARE
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FIGURE 8.6
TYPICAL OPERATING AND EMERGENCY FLARES
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FIGURE 8.7
TYPICAL QUENCH DRUM (CONDENSABLES)
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FIGURE 8.8
TYPICAL QUENCH DRUM (EMERGENCIES)
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FLARE SYSTEM
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9.0
FLARE
Flares provide the ability to dispose of flammable, toxic or corrosive gases to less
objectionable compounds by combustion or thermal oxidation. They accommodate disposal
of process relief and block valve leakages, process off-gas vents and other normal gas
streams which cannot be recycled to the refinery, on a continuing basis. During emergency
releases from the processes, a much larger quantity of gaseous material must be safely
disposed of as well.
9.1
DESIGN DISCUSSION
9.1.1
Selection of Flare Stack Location
A.
New Relief Systems
Location of the flare stack for a new relief system is dependent on the total
relief load (Btu/h, kJ/h, kcal/h), height of the flare stack, and acceptable
ground level radiation. The selection of the new flare stack location is based
on two primary criteria:
o
Available pressure drop from the relief sources to the flare tip
For low relief valve set pressures, the available pressure drop may be
low enough to limit the distance between the source of relief and the
flare tip. However, this is usually not the case. Sufficient pressure
drop is normally available to site the flare stack at any location within
the refinery plant limits.
o
Available plot space for the flare stack and safety circle
The first choice would be an open area at the extreme end of the
refinery where future expansion could be performed without impact
on any adjacent facilities, including tankage areas.
Second choice would be an open area between or adjacent to
process units. This choice would limit future flexibility because of the
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lack of plot space to expand or debottleneck the process units,
potential blockage of access to the process units by the flare facilities,
and the reduced capability to expand the safety circle with increased
future relief loads.
Third choice would be within a process area. Startup flares, for
example, are frequently located within process units because of the
relatively low thermal loads. With sufficiently low relief loads and
adequate stack height the safety circle becomes nonexistent.
B.
Revamp of Existing Relief Systems
A new flare should not be selected for revamp of an existing relief system
without first exploring the possibility of continued use of the unmodified
existing flares. This activity should consist of the following steps:
o
Check if the existing capacity is sufficient
Determine the new design loads and check the resulting relief system
hydraulics and radiant heat release. If this is acceptable then no
further engineering design needs to be performed.
o
Reduce the new design case to fit the existing system
Review the new design loads to see if the total relief loads can be
reduced to acceptable levels. This may involve additional
instrumentation, conversion of pump or compressor drivers to steam
and electrical power, modification of the electrical power distribution
system, or process modifications. Review Chapter 4 for the selection
basis for relief loads. Usually, where existing capacity is marginal for
the new loads, one or more of the larger relief loads can be reduced
in a rational manner. This additional engineering is justified to
eliminate the large capital investment required for a new flare system.
Consider performing a dynamic analysis of the process relief loads to
determine if the design loads can be reduced and/or are not in fact
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simultaneous but spread over a sufficient amount of time to take
credit for the reduction on the maximum instantaneous relief load.
o
Where flare header hydraulics are limiting, several solutions may be
available:
•
Provide a new relief header parallel to the existing relief
header(s).
•
Replace conventional relief valves with balanced bellows or
pilot operated relief valves in order to increase the allowable
pressure at the discharge of the relief valves.
•
Review the equipment with the lowest allowable relief valve
outlet pressure to determine if the equipment MAWP can be
increased. This is a specialized review that should include
vessel designers. Also, consider the existing condition of the
protected equipment with respect to corrosion.
•
Dynamic analysis, as discussed above.
Where a new flare must be added to an existing relief system, its location is
determined by the same criteria as the design of a new relief system, along with the
following considerations:
o
Locate the flare as near the relief sources as is feasible.
Extensive new piping runs with sloping headers can be very expensive.
o
Normal heat flux limits may be reconsidered by using special designs.
o
Normal heat flux limits may be exceeded by the use of radiation shielding in
operating areas. This cost is not normally acceptable for new units but could
be economical for a revamp.
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9.1.2
Flare System Data Sheet
The purpose of this chapter is to establish the design requirements for the flare and
flare stack. These requirements are to be summarized on the Flare System Data
Sheet. In order to begin this process, it is necessary to have established the
required flow rate, molecular weight and temperature of the gases to be flared.
From this information and the following procedures, the duty specification for the
flare and stack can be developed:
o
o
o
o
Design flow rate, lb/h (kg/h) for the maximum case
Molecular weight of the flare gas for the design case.
Temperature of the flare gas for the design case.
Smokeless capacity required, lb/h (kg/h), with an analysis of the flare gas to
establish steam or air requirements.
The information above is used to prepare a request for vendor quotation to supply
the flare and/or flare stack. For elevated flares, see Form E-436B for the required
format. The vendor’s proposal should be reviewed and compared with the following
calculations to ensure that suitable equipment is provided.
9.2
TYPES OF FLARES
9.2.1
Discussion
Flares are built in the following configurations:
o
o
Vertical, elevated flares are the most commonly used designs. Multi-tip
flares mounted in an elevated position are also available.
Boom-supported flares are commonly used for offshore facilities.
o
Vertical ground flares have multiple burners and are located either in
refractory lined enclosures or in open pits.
o
Horizontal ground flares are located in open pits.
o
Burn pits are open pit spargers without burners.
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FLARE
Environmental concerns must be addressed to establish the limitations on smoke,
noise, luminosity, and odor from flares. See Table 9.1 for a comparison of flare
characteristics. Hydrocarbon gases other than methane generally do not burn
smokelessly unless specially designed smokeless flares are used. These flares
normally utilize steam, forced air or turbulent mixing action to obtain improved
combustion performance and smokeless operation.
Assist fuel gas is usually required to ensure complete combustion of gases having a
lower heating value below 150 Btu/SCF (5,610 kJ/Nm3, 1,340 kcal/Nm3 ). Ammonia
has a lower heating value of 365 Btu/SCF (13,600kJ/Nm3, 3,248 kcal/Nm3 ) and may
require assist fuel gas to ensure complete combustion while minimizing nitrogen
oxide production. Gas streams containing sulfur compounds (for example, sulfur
plant tail gas) require a minimum lower heating value of 250 Btu/SCF (9,316 kJ/
Nm3, 2,225 kcal/Nm3 ) for complete combustion.
The second criterion is gas analysis. The CO2 content of the gas radically affects
the heat content required to sustain combustion.
The gas assist requirement also depends on the flare tip design and the flare gas
exit velocity. In any case, a flare vendor should be consulted about the fuel gas
requirement when a low heat content gas is to be flared.
Sealing and purging requirements are described in Section 7.0, “Relief System
Piping and Purging”.
9.2.2
Elevated Flares
The basic structural support systems for vertical, elevated flares in onshore facilities
are the self-supporting, guyed, and derrick-supported. The self-supporting stack is a
freestanding stack anchored to a base. The guyed stack is anchored by guy wires.
The diameter of the circle connecting the guy wire anchors is frequently equal to the
stack height. The derrick-supported stack is located in the center of a derrick
structure and is held to the structure by tie rods and guides.
Self-supporting stacks are preferred for heights up to 246 feet (75 meters). Derrick-
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supported stacks have been built up to 394 feet (120 meters) without significant
problems. Guyed stacks have been constructed with heights up to 591 feet (180
meters). For flare stacks higher than 591 feet (180 meters), a concrete support
structure is required.
The structural support systems for elevated flares in offshore facilities are the boomsupported and derrick-supported. The flare may be either vertical or inclined. The
boom is an inclined support structure consisting of tie rods and guides.
9.2.3
Ground Flares
Ground flares usually consist of a series of burners standpipe-mounted on
underground pipe manifolds. The burners are located either in a refractory lined
enclosure or in an open pit. Another type of ground flare is the horizontal flare which
is located in an open pit.
Hot combustion gases from an enclosed ground flare are discharged to the
atmosphere through an opening at the top of the refractory-lined enclosure. An
acoustical fence may be provided around the enclosure to reduce noise levels.
Ground level radiant heat intensities outside the enclosure are very low compared to
elevated flares. The burners can be designed to achieve smokeless flaring by
means of high velocity vortex action.
Elevated flares are preferred over enclosed ground flares due to lower costs. The
latter type is the preferred choice only if a plant is located in an area where it is
highly desirable to have a flare which is not visible to the public.
The open pit ground flare utilizes either a series of burners standpipe-mounted on
underground pipe manifolds or one large burner installed in the horizontal position
on one side of the pit.
Elevated flares are also preferred over open pit ground flares due to lower land
requirements. Open pit ground flares are seldom selected. Environmental
regulations must be checked to determine if open pit flares are allowed.
Elevated flares require the least land to install. However, if a ground flare is
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required, the ground level, staged, multitip flares will require less area.
9.2.4
Offshore Platform Flares
Three basic types of offshore flare support structure are illustrated in Figure 9.14:
o
o
o
Deck supported, vertical derrick.
Sea-bed supported (remote). This is usually connected by a bridge but may
be connected by a sea-bed pipeline.
Deck supported angle boom.
The choice of flare support structure depends on:
o
The amount of hydrocarbon to be disposed; this establishes the amount of
heat liberated by combustion. The greater the heat, the further away the
flare must be from other facilities for the same radiant heat intensity. Boom
flares are generally limited to about 459 ft (140 m).
o
The depth of the water and other environmental factors which affect the cost
of sea-bed supported structures. Horizontal bridges are typically supported
each 164 ft (50 m) but economic optimization depends on water depth.
The substitution of a subsea pipeline for a pipebridge to the flare should only be
used when liquid condensation is not possible. Condensation accumulating in the
line would cause operational high back pressures and burning liquid to be
discharged. High back pressures may also result if sea water leaks into the low
pressure line.
The risks and effects associated with accumulated liquid require dynamic simulation
studies.
Vertical flares can be a source of fire if liquid is carried over, as flowing liquid could
fall onto the platform. Vertical flares should only be used where the risk of burning
rain and bulk liquid carry-over can be minimized to an acceptable level by use of
flare tips which can properly dispose of liquid and by shutdown systems which are
initiated by high liquid level.
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Normal maintenance costs and tip replacement frequency will be economic factors
in flare type selection.
9.3
FLARE SYSTEM METALLURGY
9.3.1
Hydrocarbon Flaring
Relief streams from major process units are normally routed to the flare via the
laterals, subheaders and main flare header to the flare K.O. drum. The flare
headers and subheaders are normally swept by nitrogen or refinery fuel gas which
helps keep the piping free of vapor contaminants. Also, all the flare discharge piping
is sloped to the flare K.O. drum which collects minor liquid discharges from the
relieving devices and any condensate that may form after the relief. Based on these
design premises, carbon steel piping is normally adequate.
The flare K.O. drum is normally stress relieved carbon steel due to the fact that
contaminants such as H2S may be present in the condensed liquid.
The flare tip, including the tip assembly, usually is constructed of 310 stainless steel.
Incoloy 800H is also acceptable material for the flare tip.
9.3.2
H2S Flaring
Relief streams from refinery process units such as Sour Water Stripping, Amine
Regeneration, Sulfur Recovery and Tail Gas Treating (Claus plant) often contain
H2S. A service is considered sour if the stream has the following characteristics:
o
A wet system, containing an aqueous phase, with the total pressure greater
than 65 psia (4.5 bara, 4.6 kg/cm2 A) and partial pressure of H2S is greater
than 0.050 psia (0.0035 bara, 0.0035 kg/cm2 A), or
o
A process water stream or a mixed phase stream containing water, in which
the H2S concentration is greater than 50 ppmw.
Stress relieving is not required for flare piping in wet H2S service. However,
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stainless steel or NACE-certified valve trim may be required. Sour gas containing
ammonia (typically from sour water stripper) has a tendency to clogging flare lines
due to solidification of ammonium bisulfides when the stream temperature drops
below 158 °F (70 °C); therefore, it is recommended that heat tracing be provided to
maintain the vapor temperature above 176 °F (80°C).
An H2S flare K.O. drum, subjected to a moderate H2S service, is normally stress
relieved.
Low carbon 309 stainless steel material is recommended for all H2S flare tip.
9.4
ELEVATED FLARE SIZING
9.4.1
Discussion of Sizing Methods
o
Flare Stack
The sizing of the flare stack requires that the following items be calculated or
estimated:
•
Flare stack diameter (set by velocity and available ∆P)
•
Wind tilt (set by wind velocity and stack exit velocity)
•
Dispersion (set by stack height and wind speed)
•
Height (set by radiation intensity and ground level concentration of
emissions)
o
Flare Tip
The flare tip designs are proprietary and must be supplied by qualified
vendors.
9.4.2
Nomenclature
Unless otherwise stated, all symbols used in this section are defined as follows:
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CL
=
dimensionless lean flammability limit concentration of the relief gas
mixture.
CL
=
lean flammability limit concentration of the relief gas mixture in air,
volume fraction.
CLi
=
lean flammability limit concentration of the individual components of
the relief gas mixture in air, volume fraction.
D
=
diagonal distance from a given point, at grade (where the radiation
intensity is to be calculated or fixed) to the radiation source point, ft
(m).
d
=
flare stack tip diameter, inches.
H
=
flare stack height, ft (m).
H’
=
vertical distance from grade to the flame “radiation source point”, ft
(m).
∆h
=
net heat of combustion (lower heating value, LHV), Btu/lb (kJ/kg,
kcal/kg)
h
=
convective heat transfer coefficient over equipment surface,
Btu/ft2/h/°F (kJ/m2/h/°C, kcal/m2/h/°C).
k
=
specific heat ratio of the relief stream = CP/Cv.
K
=
radiant heat intensity, Btu/ft2/h (kJ/m2/h, kcal/m2/h)
L
=
flare flame length (straight line distance from the flare tip to the flame
tip), ft (m). (See also “SL”).
F
=
fraction of total heat released as radiant heat.
Ma
=
molecular weight of air, 28.96
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Mg
=
molecular weight of the relief gas mixture
Mach No.=
ratio of vapor velocity to sonic velocity in the vapor.
N
=
ground projected distance from the flame “radiation source point” to
the downwind point where the radiation intensity is to be calculated or
fixed, ft (m).
P
=
absolute pressure of the relief gas mixture just inside the flare tip,
psia (bara, kg/cm2 A).
∆P
=
pressure drop across the flare tip, in (mm) of water.
Qr
=
radiation heat release from the flare, Btu/h (kJ/h, kcal/h).
q
=
relief flow, ft3/s (m3/s) (actual).
R
=
dimensionless scaling parameter to account for the relative dynamic
pressure of the relief gas jet and the wind.
R
=
radius of the circle around the stack at which allowable ground level
radiant heat intensity occurs, ft (m).
Rg
=
Ideal gas constant, 10.732 psi-ft3/lbmol-°R (0.08315 bar-m3/kmol-K,
0.08479 (kg/cm2)-m3/kmol-K)
SL
=
dimensionless coordinate of the concentration CL in the axis of the
relief gas jet.
sL
=
actual flame length on the real (curved) axis of the flame, ft (m) (sL ≥
L).
Tg
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=
absolute temperature of the relief gas mixture just inside the flare tip,
°R (°K).
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Ta
=
absolute ambient air temperature, °R (°K).
Ts
=
hot spot temperature on equipment surface, °R (°K).
Ua
=
horizontal wind velocity, ft/s (m/s).
Ug
=
vertical velocity of the relief gas mixture exiting the flare tip, ft/s (m/s).
Us
=
sonic velocity in a vapor, ft/s (m/s).
V
=
relief flow, ft3/h (m3/h) (actual).
W
=
relief flow, lb/h (kg/h)
XL
=
dimensionless downwind coordinate of the flare tip.
xc
=
horizontal coordinate of the “radiation source point”, ft (m).
xL
=
horizontal coordinate of the flame tip, ft (m).
yc
=
vertical coordinate of the “radiation source point”, ft (m).
yi
=
mole fraction of the individual components of the relief gas mixture.
zL
=
dimensionless rise of the flame tip over the flare tip.
zL
=
vertical coordinate of the flame tip, ft (m).
εs
=
surface emissivity
θ
=
angle of flame tilt from vertical, due to wind velocity, degrees.
ρa
=
density of the air, lb/ft3 (kg/m3).
ρg
=
density of relief gas mixture before exiting the flare tip, lb/ft3 (kg/m3).
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σ
=
Stefan-Boltzman constant, 1.712 x 10-9 Btu/ft2/h/°R4 (2.05 x 10-7
kJ/m2/h/°K4, 4.89 x 10-8 kcal/m2/h/°K4)
τ
9.4.3
=
fraction of radiant-heat transmitted
Stack Diameter
The flare stack diameter is generally sized on a velocity basis. However, the
pressure drops through the stack, tip, and baffle or labyrinth seal (if any) must be
checked. Although pressure drops as high as 12 in (300 mm) of water (2.8 psi, 0.2
bar, 0.2 kg/cm2 ) have been satisfactorily used at most types of flare tips, a flare
vendor should be consulted to finalize a satisfactory pressure drop. The mixing
action type of flare tip requires a minimum pressure of 10 psi (0.69 bar, 0.70 kg/cm2).
Refer to Section 9.6, “Smokeless Flaring”, for further information on pressure
requirements.
The following criteria can be used to estimate the required diameter:
o
For the maximum instantaneous relief flow to a new flare, the velocity at the
tip should not exceed a Mach No. of 0.5. Sizing is primarily governed by
allowable backpressure with modern flares which can stabilize flames at a
higher Mach No. For existing flares, do not exceed a Mach No. of 0.7.
o
When flaring gas with a low heating value, the tip velocity may need to be
decreased to as low as 59 ft/s (18 m/s) depending on the tip design. This
will result in a large diameter tip. The riser and flare line velocities should
follow the normal criteria.
The gas sonic velocity is calculated by:
Us = 68.0 [(k Tg) / Mg ]0.5
Us = 91.2 [(k Tg) / Mg ]0.5
9.1 (English)
9.1 (Metric)
The equation used to estimate the pressure drop for flare tips other than the mixing
action type is presented below. The pressure drop across a baffle or labyrinth type
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seal may also be estimated by this formula. See Section 7.0, “Relief System Piping
and Purging” for further information on seals.
∆P = 7.37 x 10-6 ( Tg /Mg) ( W 2 /d4), in H2O
9.2 (English)
∆P = 1.64 x 10 ( Tg /Mg) ( W /d ), mm H2O
9.2 (Metric)
-3
9.4.4
2
4
Stack Height
The flare stack height is based upon the radiant heat intensity generated by the
flame and upon the available clear area around the stack. If toxic or corrosive
pollutants are present in the relief stream, the maximum pollutant concentration
level, the stack height, and its location must be checked to determine what potential
toxic and/or environmental hazard might occur if the flare pilot flame is lost.
Maximum flare height is also restricted by aircraft clearance requirements.
Selection of allowable radiant heat intensities for use in flare stack height sizing
involves many factors such as probability of maximum releases, duration and
directions of release, wind velocity, number of stacks, and need for personnel
activity in exposed areas. From the standpoint of human safety, allowable ground
level radiant heat intensities (not including solar radiation) when the maximum
instantaneous relief flow to the flare occurs are as follows:
o
500 Btu/ft2h (5,677 kJ/m2/h, 1,356 kcal/m2/h) (not including solar radiation) in
operating areas where operators, wearing normal clothing, are required to
perform their duties without the benefit of radiant shielding for long term flare
operation at the maximum continuous relief load.
o
1,500 Btu/ft2/h (17,036 kJ/m2/h, 4,069 kcal/m2/h) in operating areas where
operators wearing normal clothing are likely to perform their duties and
where general area radiant shielding is available.
o
3,000 Btu/ft2/h (34,068 kJ/m2/h, 8,137 kcal/m2/h) where access is limited and
shelter is available. Exposure should be limited to a few seconds.
Allowable personnel exposure times are: 5 minutes for 1,500 Btu/ft2/h (17,036
kJ/m2/h, 4,069 kcal/m2/h) and several seconds for 3,000 Btu/ft2/h (34,068 kJ/m2/h,
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8,137 kcal/m2/h).
The radiation heat release from a flare is considerably less than the total heat
release because most of the heat is lost by convection. The radiation heat release
from the flare is calculated by the following equation:
Qr = FW∆h
9.3
The factors that affect the value of F are very complex, and depend more on the
composition of the combustion products (CO2, H2O and any remaining C or CO) than
the composition of the gas being flared. They also depend (according to the
literature) on the flaring rate, wind velocity, and atmospheric condition.
The stack height calculation is based upon reducing the flare to an equivalent point
source and taking into account the effect of wind on flame tilt. The geometry of the
stack and flame, as affected by the wind, is illustrated in Figure 9.3.
The equations used to calculate the stack height and the radiation intensity at grade
are those proposed by T.A. Brzustowski and E.C. Sommers (B&S Method), as
described in API RP 521. Other sizing methods which are sometimes referenced
include the API Simple Method and the Kent Method. The API Simple Method is
discussed in API RP 521 and the Kent Method is presented in an article by G.R.
Kent (8). These methods differ primarily in the approach used to calculate the flare
center. In general, the B&S calculation method provides radiation levels at grade
which are intermediate between the other methods, with the API Simple Method
providing the most conservative (highest) estimate of radiation levels.
The general steps to be followed for calculating the flare stack height:
a)
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Establish a flare area (circle) with piping design group, where personal
access shall be limited. This circle shall not include process plot areas,
combustible material storage tanks, utility areas and any other facilities
where personal access is required for daily activities. Flare K.O. drum and
seal drums may be located inside this circle, if proper shelter is provided near
the K.O. drums and/or seal drums. (Note: US insurance industry guidelines
suggest a minimum spacing of 330 ft (100 m) for flare stacks less than 82 ft
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(25 m) in height and a minimum spacing of 213 ft (65 m) for taller flare stacks
from general refinery facilities.)
b)
Calculate the stack height so that the maximum ground level radiant heat
intensity does not exceed 1,500 Btu/ft2/h (17,036 kJ/m2/h, 4,069 kcal/m2/h)
at the edge of the flare area circle, based upon the maximum instantaneous
relief flow to the flare and the design wind velocity. If free personal access is
required at any flare area, the stack height shall be sized not to 1,500
Btu/ft2/h (17,036 kJ/m2/h, 4,069 kcal/m2/h) at any ground level.
c)
Using the height and radius calculated in the first step and the normal wind
velocity, check that the radiant heat intensity does not exceed 500 Btu/ft2h
(5,677 kJ/m2/h, 1,356 kcal/m2/h) at a distance from the base of the stack
equal to this radius when the maximum continuous relief flow to the flare
occurs.
d)
Using the procedures described in Section 9.4.6, “Equipment Surface
Temperature”, check that the surface temperatures of process equipment
and storage tanks do not exceed 302 to 392 °F (150 to 200 °C) when the
maximum relief flow occurs. Keep in mind that long term continuous flaring
may create expansion problems for equipment and pipe at elevated ground
level temperatures which exceed normally expected continuous values.
e)
Structural limitations are discussed in Section 9.2.2 “Elevated Flares”.
A sample calculation of radiation levels from an existing flare is provided in Appendix
B-6 and the spreadsheet program for this calculation is provided in Appendix D.
9.4.4.1 Data Required
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o
Gas flow rate; maximum instantaneous (peak) relief flow rate, and
maximum continuous (normal) flow rate.
o
Approximate composition of the gas to be flared. If there is a
substantial difference between the composition of the gas during the
maximum instantaneous relief and that of the maximum continuous
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relief, both compositions are needed. If composition is not available,
approximation can be done based on molecular weight.
o
Temperature of the gas to be flared.
o
Low (net) heating value (LHV) of the gas to be flared. LHV can be
estimated from molecular weight assuming that the flare gas is all
hydrocarbon and/or hydrogen.
o
Wind velocity pattern
The normal wind velocity is taken as the velocity which is exceeded
50 percent of the time. If no data is available, use 10 ft/s (3 m/s) for
the normal wind velocity.
The design wind velocity is taken as the lower of 29 ft/s (8.9 m/s) or
the velocity exceeded 5 percent of the time. If no data is available,
assume 29 ft/s (8.9 m/s).
o
Stack diameter (previously calculated).
9.4.4.2 Procedure for Stack Height Calculation
1)
Calculate the molecular weight of the gas based on its composition.
2)
Calculate the gas density by:
ρg = (Mg P) / (Rg / Tg)
3)
Calculate the volumetric gas flow rate by:
V = W /ρg
4)
9.5
Based on the tip diameter, calculate the gas exit velocity by:
Ug = 0.0509 V/d2 (for d in inches)
Ug = 0.549 V/d2 (for d in inches)
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9.4
9.6 (English)
9.6 (Metric)
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5)
Calculate the average lean flammability limit of the gas mixture, CL,
by using the lean flammability limit of the individual gas components,
given on Table 9.2, and the mole fractions of these components.
CL = [1/=Σ(Yi/CLi)]
6)
Calculate dimensionless lean flammability limit concentration (volume
fraction) of the gas being fired, CL.
CL = CL (Ug / Ua) (Mg / Ma )
7)
9.7
9.8
Calculate the dimensionless coordinate of the concentration CL on
the axis of the flame SL , and from it, the downwind coordinate XL.
This latter coordinate is then identified with the downwind location of
the flame tip (in the “X” axis).
i)
ii)
If CL ≤ 0.5
a)
calculate SL = 2.04 / (CL )1.03
9.9
b)
calculate XL = SL - 1.65
9.10
If CL > 0.5
a)
Calculate SL = =2.5 / (CL )0.625
b)
If SL > 2.35 calculate
XL =  SL - 1.65
c)
9.11
9.12
If SL ≤ 2.35 findXL graphically using the attached
Figure 9.2
8)
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Calculate the dimensionless rise ZL of the flame tip over the flare tip:
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ZL = 2.05 (XL)0.28
9)
9.13
Calculate the dimensionless scaling parameter R which accounts for
the relative dynamic pressures of the jet of gases (flame) and the
wind.
R = (Ug / Ua) (ρg / ρa)0.5
10)
11)
12)
13)
Calculate the full-scale coordinates (actual coordinates) of the flame
tip, meters.
zL = (ZL dR) / 39.4
9.15
xL = (XL d R) / 39.4
9.16
Calculate the coordinates of the “radiation source point”:
xc = xL / 2
9.17
yc = 0.82 zL
9.18
If required, estimate the flame tilt angle from the vertical by:
θ = tan -1 (xc /yc )
9.19
or tan -1 (xL / zL)
9.20
If required, estimate the straight flame length (distance between the
flare tip and the flame tip) by:
L = [ (zL)2 + (xL)2 ] 0.5
14)
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9.14
9.21
If the actual curved length of the flame is required, estimate this value
by:
9.22
sL = (SL ) ( d / 39.4 ) (R )
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At this point, calculate either the stack height (H) for a fixed maximum
radiation intensity at a desired downwind distance (R) or calculate the
radiation intensity (K) at any downwind distance for a given stack height.
Referring to Figure 9.1:
H’ = H + yc
9.23
N = R - xc
9.24
D = [ H’2 + N2 ] 0.5
9.25
The radiation intensity (K) at any given point at grade is a function of the heat
release (Qr) and the diagonal distance (D) from that point to the flame
“radiation source point”. See Figure 9.1.
The heat release, Qr , was previously calculated by:
Qr = (W) (F) (∆h)
(9.3)
The relation between the diagonal distance (D) and the radiation intensity (K)
is given by:
D = [ Qrτ / (4πK) ] 0.5 = [ (W) (F) (τ) (∆h) / 4πk ]0.5
9.26
The factor “τ” accounts for radiant heat absorbed by the atmosphere and
typically is related to the amount of moisture in the air and flared gas.
See Section 9.4.5 for appropriate values for F. Since “D” is the diagonal
distance that depends on the stack height (H) and the horizontal distance
between the base of the stack and the point at which the radiation intensity is
measured (R), it is obvious that either one of these parameters can be
calculated once the other is fixed.
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9.4.5
Radiation Considerations
9.4.5.1
Fraction of Heat Radiated (F)
The F-factors given below are typical of the values recommended by
the major flare vendors. These choices are based on consideration
of overall safety. The less conservative value for maximum flaring is
balanced by the other more conservative assumptions for peak flare
loads. The more conservative value for continuous flaring is
warranted by its probability and frequency.
F ==
9.4.5.2
•
Maximum flaring for 10 minutes or less
0.15
•
Continuous flaring
0.25
Fraction of Heat Intensity Transmitted
A value of 1.0 should be used for the fraction of heat intensity
transmitted in Equation 9.26, when using the values of F provided in
Section 9.4.5.1, consistent with the basis used in selecting the
values for F.
9.4.5.3
Heat Absorbed by Adjacent Equipment
For equipment protection, the heat intensity on structures must be
limited to 5,000 Btu/ft2 h (5,677 kJ/m2h, 13,562 kcal/m2h) and in
areas where operators are not likely to be performing duties and
where adjacent shelter is available.
9.4.6
Equipment Surface Temperature
The surface temperatures of process equipment and storage tanks, when the
maximum relief flow to the flare occurs, can be estimated using the following
equations:
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K = σ ( Ts4 - Ta4 ) + =(h/=εs )( Ts - Ta )
o
For surfaces sheltered from the wind:
h = 0.214 (Ts - Ta ) 1/3
h = 5.32 (Ts - Ta ) 1/3
h = 1.27 (Ts - Ta ) 1/3
o
9.27
9.28 (English)
9.28 (Metric)
9.28 (Metric)
For surfaces exposed to the wind:
h = 2.05 Btu/ft2/h/°F
h = 42 kJ/m2/h/°C
h = 10 kcal/m2/h/°C
for wind velocity up to 7 ft/s (2 m/s)
9.29 (English)
9.29 (Metric)
9.29 (Metric)
9.30 (English)
h = 4.1 Btu/ft2/h/°F
2
h = 84 kJ/m /h/°C
9.30 (Metric)
2
h = 20 kcal/m /h/°C
9.30 (Metric)
for wind velocity between 7 ft/s and 30 ft/s (2 m/s and 9 m/s)
9.31 (English)
h = 6.2 Btul/ft2/h/°F
2
h = 126 kJ/m /h/°C
9.31 (Metric)
2
h = 30 kcal/m /h/°C
9.31 (Metric)
for wind velocity greater than 30 ft/s (9 m/s)
These relationships were derived assuming steady state conditions when
heat loss into the equipment and its contents becomes negligible. Therefore,
for short term flaring, the hot spot surface temperatures predicted by these
relationships are conservative. Exposure times required to reach these
temperatures can be estimated using Figure 9.3.
The following example will illustrate the estimation of equipment surface
temperature.
Example: Estimate the surface temperature on cast iron equipment which is
receiving radiation from the flare at 498 Btu/ft2/h. The air temperature is 61
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°F and wind velocity is 19 ft/s.
K = 498 Btu/ft2/h
Ta = 61 + 460 = 521 °R
From Table 9.3, εs for Cast Iron = 0.8
h = 4.1 Btu/ft2/h/°F from wind velocity of 19 ft/s
Rearrange the equation
σTs4 + (h/εs) Ts = K + σTa4 + (h/εs) Ta
9.32
= 498 + 1.71 x 10-9 (521)4 + (4.1/0.8) (521)
= 3294
The equipment surface temperature Ts is calculated by trial and error.
To obtain an initial estimate of Ts, neglect σTs4 from the above equation and
solve for Ts.
Ts (guess) = 3294 x ( εs /h) = (3294) ( 0.8/4.1) = 643 °R
Guess different values of Ts in the equation and compute the left-hand side of
the equation. When the left-hand side equals the right hand side of the
equation, the correct surface temperature Ts is obtained.
Ts = 599°R = 140 °F
9.5
GROUND FLARE SIZING
9.5.1
Enclosed Ground Flares
Enclosed ground flares are the most practical ground flare type for industries located
in the city areas. While they are expensive, they provide a means of smokeless
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combustion without a flame visible to the surrounding area.
Enclosed ground flares are limited to a maximum relief load capacity of about 100
t/h. Due to this capacity limitation and the high cost of enclosed ground flares, they
are most often used to handle continuous flaring loads and minor upsets, with major
emergency loads being diverted to an elevated flare.
A typical enclosed ground flare installation is shown in Figure 9.4. Vapor loads
normally pass through the low pressure seal and to the enclosed ground flare where
multiple small burners are used to combust the vapors. These burners are divided
into groups which are activated (staged) via control valves, with additional stages
opening as header pressure increases. Since only the first group (stage) of burners
is operating continuously, purge requirements for enclosed ground flares are based
on maintaining a required flow through those burners only. When a major upset
occurs, pressure rises in the header to the point where flow passes through the
liquid seal to the emergency elevated flare which is sized to handle the worst case
flaring contingency.
The enclosed ground flare design is based upon the ground level radiant heat
intensity outside the enclosure being less than solar radiation. Therefore, this type
of flare can be located in areas where operators wearing normal clothing are
required to perform their duties without the benefit of radiant shielding. The relief
gas pressure required at the burner is similar to that required for elevated flares. If
steam assist is used to provide smokeless flaring, gas pressure near atmospheric
can be used. A minimum pressure of about 6 psig (0.4 barg, 0.4 kg/cm2 G) is
required to achieve smokeless flaring without steam assist.
The height of the flare enclosure is based upon protecting personnel from the hot
combustion gases discharged to the atmosphere as well as from the burner flames.
A preliminary estimate can be made using the following criteria:
o
The minimum height is 25 ft (8 m) to allow for burner height and flame length.
o
The top of the enclosure is a minimum of 10 ft (3 m) above any building or
platform within a radius of 25 ft (8 m).
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The following criteria can be used to make a preliminary estimate of the flare
dimensions (for a cylindrical design) and the number of burners:
o
Inside volume of flare shell (enclosure) = 1590 ft3/t/h (45 m3/t/h) gas flared
o
L/D for cylindrical flare shell = 3/1
o
Enclosure (fence) diameter = 2 x shell diameter
o
Maximum gas flow to each burner = 17700 SCFH (500 Nm3/h)
A vendor must be consulted to finalize the height of the flare enclosure, the
size and layout of the burners, and the relief gas pressure required at the
burners.
9.5.2
Open Pit Ground Flares
The location of an open pit ground flare is determined as follows:
o
Calculate the distance that personnel and equipment must be from the wall
around the open pit based upon the radiant heat intensity criteria used for
elevated flares and based upon the following formula:
Distance = [WF∆h / 4πK] 0.5
o
9.33
Process equipment and storage tanks shall be located a minimum distance
of 200 ft (61 meters) from the wall around the pit to prevent fire hazards from
the flame or flames.
The preliminary burner size and layout for a vertical open pit ground flare is
estimated according to the procedures used for the enclosed ground flare. The relief
gas pressure required at the burner is the same as for the enclosed ground flare.
The minimum pit depth is 25 ft (8 m) to allow for burner height and flame length.
The burner diameter and flame length for a horizontal flare handling gases are
estimated according to the methods used for elevated flares. The length of the pit is
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equal to the flame length while the width is one-half of the length. The depth is
based upon the burner diameter plus allowances of 4 ft (1.2 m) above the burner
and 2 ft (0.6 m) below the burner.
An open pit for a horizontal flare handling liquids is sized to provide a liquid holdup
volume based upon the maximum flow and the maximum duration of this flow. The
burner diameter is generally smaller than the size of the liquid line to the burner.
The liquid pressure required at the burner is typically 20 psig (1.4 barg, 1.4 kg/cm2
G). The depth is based upon the burner diameter plus allowances of 4 ft (1.2 m)
above the burner and 2 ft (0.6 m) below the burner.
The horizontal flare is flush-mounted in the pit wall so that the outside of the tip is not
exposed to an open flame. The pit floor is sloped away from a horizontal flare
handling liquids.
A vendor must be consulted to finalize the dimensions of the pit, the size and layout
of the burner or burners, and the relief stream pressure required at the burner or
burners. Open pit ground flares require a significant plot area which in many
locations makes them impractical.
9.5.3
Burn Pit
A burn pit (shown in Figure 9.5) is not acceptable for frequent use because of the
degree of smoking caused by lack of burners to provide an appropriate air/fuel
mixture. It may sometimes be considered for handling releases with a 20 to 40 year
frequency, if adequate plot space is available. A vendor must be consulted to
finalize the dimensions of the pit, the layout of the distribution, and the pressure
required at the distribution inlet.
9.6
SMOKELESS FLARING
Hydrocarbon gases other than methane generally do not burn smokelessly unless specially
designed flares are used. These flares utilize steam, forced air, turbulent mixing action,
water spray, or fuel gas to obtain improved combustion performance and smokeless
operation. The water spray and fuel gas smokeless flares are the least preferred types.
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Local environmental regulations must be checked to determine the limitations on smoke
from a flare. In most cases, smokeless operation is not required at the maximum
instantaneous relief load used to size the flare.
9.6.1
Smokeless Flaring Requirements
The required smokeless flaring capacity is usually based on the maximum
continuous relief load which can exceed five minutes duration during any 2
consecutive hours.
An acceptable initial estimate for the smokeless flaring rate would be 10 percent of
the design relief. This estimate must be confirmed during relief system design
development. If larger rates are anticipated for process unit startups, shutdowns or
depressuring operations, then the larger requirement must be weighed against
available steam. During emergency relief, the steam may be better utilized in
turbine drivers to minimize both the process upsets and the resulting total relief load.
Ultimate smokeless capacity is determined by vapor properties, tip design and,
where applicable, steam supply hydraulics.
9.6.2
Steam Injection
Elevated flares generally utilize steam for smokeless operation. A typical steam
assisted flare tip is shown in Figure 9.9. Steam requirements for smokeless
operation of this tip or equivalent are determined from Figure 9.7. However, most of
the flare vendors have improved the steam injection nozzles and air inducing
method for elevated flare tips recently, achieving significant reduction of steam
requirements (about 20 to 30% of that obtained from Figure 9.7). The flare system
process engineer should consult with vendors for the latest information in the steam
requirement. The minimum steam pressure required at the steam inlet to the flare
tip is 50 psi (3.4 bar, 3.5 kg/cm2). Lower pressures cannot be used unless agreed to
by the flare tip vendor.
9.6.3
Air Assisted Flaring
Forced air is the preferred alternative to steam for smokeless operation of elevated
flares where steam is not available. The forced air is introduced in an outer stack by
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means of an air blower. The flare gas passes through an inner stack and mixes with
the air at the flare tip. The air blower requires approximately 1.6 hp per 2200 lb/h
(1.2 kW per 1,000 kg/h) of flare gas being burned smokelessly.
A typical air-assisted flare system is shown in Figure 9.6.
The horizontal open pit flare normally utilizes forced air for smokeless operation.
With the air-assisted flare, the amount of low pressure air required for smokeless
operation is only a fraction of the stoichiometric combustion air required. For
paraffins, the amount is generally 30 percent and for olefins 40 percent of the
stoichiometric combustion requirement.
The formula below presents the stoichiometric combustion equation for methane.
CH4 + 2O2 + 7.5 N2 → CO2 + 2H2O + 7.5 N2
9.34
The amount of air required for stoichiometric combustion of typical hydrocarbon
gases is given in Table 9.4.
Note that even though the moles of air required per mole of hydrocarbon gas
increase as the molecular weight of the gas increases (within the same category) the
kgmoles of air required per million kcal of heat release are fairly constant. With the
exception of hydrogen sulfide, the amount of air required for stoichiometric
combustion can be conservatively calculated at 0.796 lbmoles of air per million Btu
heat release (0.755 kgmoles/106 kJ, 3.16 kgmoles/106 kcal).
This simplifies the calculation for determining the quantity of low pressure air
required for a given situation. For example, if a hydrocarbon stream containing
about a dozen components is to be smokelessly flared with an air-assisted flare
system; instead of having to calculate the stoichiometric air required for each
component, multiplying that by the flow rate of each component and summing the
products to obtain a total, one can simply multiply the total net heat release in
millions of Btu (kJ, kcal) per hour by 0.796 (0.755, 3.16). A further simplifying
assumption can be made that most hydrocarbon gases have a lower heating value
of approximately 19,980 Btu/lb (46,500 kJ/kg, 11,100 kcal/kg). Consequently, the
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following formula is used to determine the amount of stoichiometric air required
(SAR) for a flare relief load:
SAR = 0.796 x 10-6 (W) (∆h)
-6
Where:
With
9.35 (English)
SAR = 0.755 x 10 (W) (∆h)
9.35 (Metric)
SAR = 3.16 x 10-6 (W) (∆h)
9.35 (Metric)
SAR = stoichiometric air required, lbmoles/h (kgmoles/h)
W
= gas to be flared, lb/h (kg/h)
∆h
= net heat of combustion (LHV), Btu/lb (kJ/kg, kcal/kg)
∆h
= 19,980 Btu/lb (46,500 kJ/kg, 11,100 kcal/kg) which reduces to:
SAR = 0.0159 W
SAR = 0.0351 W
SAR = 0.0351 W
9.36 (English)
9.36 (Metric)
9.36 (Metric)
The required blower horsepower to obtain smokeless operation can be estimated
from the following:
Where:
HP
=
(16.8 x 10-3 ) (a ) / n
HP
=
blower power, kW
a
=
kgmoles/h of low pressure air required for smokeless
operation
n
=
efficiency
9.37
-Use 0.5 if no better value is available
Note: The amount of air required for smokeless operation is “a”; not the amount of
air required for stoichiometric combustion (SAR).
The blower for an air-assisted flare system is usually provided with a two-speed
motor. With low flare gas rates, the low motor speed can be used. At a lower motor
speed, smokeless operation is still achieved, and significant power savings is
realized because only 1/8 of the connected power is used at the low speed. A
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sensor in the flare header can be used to switch the motor to high speed upon
greater flare gas flow rates.
9.6.4
Miscellaneous
o
Turbulent Mixing
The enclosed ground flare and the vertical open pit flare are equipped with
burners designed to achieve smokeless operation by means of high velocity
through small nozzles. Therefore, assist media for smokeless flaring are
rarely required. The relief gas pressure required at the burners is typically 5
psig (0.34 barg, 0.35 kg/cm2 G).
There are elevated flares which utilize turbulent mixing action to achieve
smokeless operation without assist media. These flares are primarily used in
oil production, gas gathering and other facilities requiring smokeless flaring at
the maximum relief gas flow. The minimum gas pressure at the flare tip inlet
is 10 psig (0.69 barg, 0.70 kg/cm2 G) for smokeless flaring. When a
turndown ratio of 2:1 is specified for this operation, the required pressures
are typically 30 psig (2.06 barg, 2.11 kg/cm2 G) for the maximum flow and 10
psig (0.69 barg, 0.70 kg/cm2 G) for the lower flow. Multiple flares and/or
higher maximum operating pressures are usually necessary for greater
turndown capabilities. One type of elevated flare utilizes a variable slot
mechanism to give a turndown ratio up to 20:1 with pressures of 30 psig
(2.06 barg, 2.11 kg/cm2 G) for the maximum flow and 10 psig (0.69 barg,
0.70 kg/cm2 G) for the lower flow. However, a backup flare or flares must be
provided because the mechanisms may fail.
o
Staged Flaring
A new design of mixing action type flares is the John Zink Linear Relief Gas
Oxidizer (LRGO). They are staged to handle several levels of gas releases
with 10 (0.69, 0.70) to 15 psig (1.03 barg, 1.05 kg/cm2 G) pressure. The
staging is accomplished by individual control valves to each incremental
bank of tips. The control valves have a rupture disc in parallel with each
valve so that the flare header may not be blocked in. Only the low level
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(leakage) bank of tips remains open continuously.
o
Water Spray
o
Water spray smokeless flares require a minimum pressure of 60 psig (4.1
barg, 4.2 kg/cm2 G) at the water inlet to the flare tip. Water consumption for
smokeless flaring is approximately five times the steam consumption for
smokeless flaring. Water spray is not used in the modern flare system.
Fuel Gas
Fuel gas smokeless flares require clean burning gas at a minimum pressure
of 60 psig (4.1 barg, 4.2 kg/cm2 G). Fuel gas consumption for smokeless
flaring is approximately equal to steam consumption for smokeless flaring.
Fuel gas is used when steam is unavailable.
9.6.5
Smokeless Flaring Control
If a proprietary smokeless flare is purchased, the manufacturer should be consulted
about the minimum necessary steam rate. The rate of steam admission can be
controlled either automatically or manually. Manual control usually involves remote
operation of a steam valve by operating personnel assigned to a unit from which the
flare is readily visible. This methodology is satisfactory if short-term smoking can be
tolerated when a sudden increase in flaring occurs. With a manual arrangement,
close supervision is required to ensure that the steam flow is reduced following the
correction of an upset; otherwise, operating costs can be excessive.
Automatic control of the steam-to-hydrocarbon ratio requires instrumentation that
measures pressure or flow in the flare system or radiation of the flame. At low
flaring rates, the fluctuation in either pressure or flow are so minute that very
sensitive instruments are required to provide sufficient steam for smokeless
combustion while at the same time avoiding waste. Controls should therefore be
carefully sized, precisely adjusted, and properly maintained to obtain satisfactory
operation.
When adequate pressure is available in the flare system over the range established
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for smokeless burning, the energy available from the pressure reduction can be
used to inspirate air at a venturi section, thus providing sufficient air in combination
with good mixing to produce smoke-free combustion. Such flares operate as a
premix burner. Because flare lines normally operate at very low pressure, these
systems have limited application. Such installations also have the disadvantage of
producing relatively high noise levels.
9.7
FLARE TIP DESIGN OPTIONS
A flare tip is a large burner or arrangement of smaller burners fitted at the end of the flare
riser. The essential requirement of the tip is to safely ignite and burn all discharges of
flammable gas. The flare tip designs are being improved continuously. Therefore, the
process engineer should obtain the latest information from flare vendors for the specific
project.
The main requirements of a pilot burner system are:
o
To maintain flame under all operating conditions
o
To re-ignite rapidly after accidental extinguishment (flame out).
In order to assure reliable pilot burner operation, it is normal practice to provide remotely
operated pilot burner ignition and monitoring systems.
9.7.1
Flare Tip Characteristics
Ideally, a flare tip has the following characteristics:
a)
Wide flow range (zero to maximum emergency).
b)
Ability to handle varying compositions.
c)
Reliable operation in high winds.
d)
Low pressure drop.
e)
Low noise levels.
f)
Minimum smoke emission.
g)
Minimum utility requirements (steam, power, high pressure gas, etc.).
h)
Low weight.
i)
Long life.
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j)
k)
9.7.2
High corrosion resistance to gas and atmosphere.
Minimum maintenance requirements.
Open Pipe Flare Tip
This is the simplest type of flare tip (Figure 9.8) consisting of a cylindrical fabrication
in stainless or high alloy steel from which the flared gas exits. It is normally provided
with an inlet flange for connection to the flare riser pipes. This type of flare tip can
be used in areas of the world where smoking is allowable or used for gases that do
not smoke such as methane, hydrogen, hydrogen sulfide and carbon monoxide.
The main features of this type of tip are:
o
Continuously operating pilot burners are attached to the periphery of the flare
tip to ignite the flared gas.
o
A perforated, slatted or cylindrical windshield is attached to the exterior of the
flare tip to prevent or resist attachment to the flame to the outside of the flare
tip in the low pressure regions which form to the sides of and downwind of a
cylinder subjected to transverse air flow.
o
A flame retention device, usually castellation or an annular nozzle attached
to the outlet of the flame tip. This tends to reduce flame lift off at high exit
velocities and reduces the possibilities of flame blow out.
o
Sometimes an inner refractory lining is provided with large diameter flare tips
to minimize the thermal degradation of the tip caused by burnback.
9.7.2.1
Open Pipe Flare Tip with Steam Injection
Smoke emission and radiation levels are reduced by the addition of steam
to conventional pipe flare tips (See Figure 9.9). Steam assisted elevated
flare tips are more typically used for refinery applications.
9.7.2.2
Open Pipe Flare Tip with High Pressure Gas Injection
Usually the gas flared originates from high as well as low pressure sources.
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The kinetic energy in this high pressure gas can be used to entrain additional
air by piping it separately to the flare tip and injecting it into the flame via a
manifold and jets in a similar fashion to steam injection systems.
The entrainment of air by the high pressure gas jets improves the air/gas
mixing at the root of the flame. This is believed to prevent the formation of
an unburnt core of gas, which would otherwise partially crack to form free
carbon with a resulting increase in smoke production and radiation.
The arrangement of the flare tip is similar to that shown on Fig. 9.10.
9.7.2.3 Open Pipe Flare Tip with Both High Pressure Gas Injection and Water or
Steam Injection
High pressure gas injection and water or steam injection can be combined in
a single flare tip to further reduce thermal radiation levels.
9.7.3
Forced Draft (Air Assisted) Flare Tips
In principle, stoichiometric combustion of the relief gas can be achieved by installing
blowers to deliver the required air flow. Since a substantial amount of air is
entrained by the pipe flare itself, only a small proportion of the stoichiometric air
requirements needs to be added to reduce smoke emissions (see Figure 9.11). The
air flow rates required are, however, very large; therefore forced draft flares are
limited in the relief duty which can be handled, because of air mixing and stack
diameter limitations.
9.7.4
Multi Tip Flares
An open pipe flare entrains air by viscous friction between the boundary of the gas
stream and the relatively stagnant atmosphere. Air entrainment is, therefore, a
function of tip perimeter length; a large diameter pipe flare tip entrains less air per
unit of gas flow than a similar, small diameter tip. The use of multiple, small
diameter burners provides improved mixing by splitting the gas and air into smaller
units. This improved mixing reduces the mixing path and results in a shorter,
broader, flame.
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The more efficient combustion of multiple tips results in lower smoke emissions and
radiation levels than a comparable single tip. Also, the shorter flame length may be
shielded. The use of multi-tip flares is thus a method of reducing platform radiation
levels and the length of the flare supporting structure.
9.7.5
Coanda Flare Tip
The skin adhesion (Coanda) effect enables high velocity gas streams to be induced
to flow in curved paths without having to constrain both sides of the gas film within
solid walls as in a conventional nozzle. In a coanda burner the effect is used to
allow the gas film to be in contact with the atmosphere; this results in the tip having a
high ratio of perimeter to area and hence the advantages of improved air
entrainment. This effect has been applied on high pressure gas flare tips.
o
Internal Coanda Flare Tips
One type of flare tip (Fig. 9.12) consists of a multitude of small internal
coanda nozzles. Each nozzle is closely coupled to its neighbors to form a
continuous manifold which is fed with high pressure gas. The nozzles are
arranged to produce a hollow cylinder of gas which entrains air as it rises
vertically. A flame is established above the matrix of burners. Alternatively,
large single units may be supported from a common manifold.
o
External Coanda Flare Tips
This type of flare (Fig. 9.13) uses the coanda effect to produce a jet of high
pressure gas passing over the surface of a hollow tulip shaped tip which may
have a low pressure gas flowing through its core. The high pressure stream
follows the curvature of the tulip and entrains and premixes air prior to
combustion while the low pressure gas burns in the high pressure gas flame.
Coanda burners have the benefits of reduced smoke generation, ease of
ignition (even under sudden venting), the ability to burn gas and entrained
liquid mixtures and reduced radiation levels. The operation of this type of
burner does, however, require high pressure gas. In addition, higher noise
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levels are experienced than with similar sized pipe flares. There is normally
only a limited optimum flow range.
9.7.6
High Velocity Tips
The use of very high and sonic velocity gas jets at the flare tip improves air
entrainment compared to similar conventional pipe flares. The use of multiple jets
and/or a low pressure jet (Fig. 9.10) ensures stabilization of flame under these high
velocities. The improved air entrainment as a result of the turbulence in the high
velocity gas jet reduces smoke formation and thermal radiation.
9.8
NOISE AND ENVIRONMENTAL
9.8.1
Noise Standards
Noise limitations for a flare shall be established in the project design basis.
Typically, the flare noise shall be less than 85 dBa at the base of the flare. Also,
noise level at the plant boundary should not exceed 75 dBA at the maximum
smokeless flaring rate.
9.8.2
Noise Discussion
Operation of flares results in noise during upset or normal operating conditions. A
major source of noise is the continuous injection of steam at the flare tip for
smokeless operation and tip cooling. The combustion process also produces noise.
The magnitude of the noise varies with the relief rate, gas composition, type and tip
design. The selected flare vendor usually makes sure that the noise level is lower
than the specified level. The required design parameters are specified in the
mechanical data sheets.
Once the noise level has been established at any point, noise levels at other
locations can be calculated by using the methods given in API RP 521 (Third
Edition), paragraph 5.4.4.3.
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9.8.3
Environmental
The environmental regulations and codes applicable for the flare design differ by
plant location. The lead process engineer shall identify the local environmental
requirements with clients and include the specific requirements in the project design
criteria.
9.9.
FLARE IGNITION
9.9.1
Discussion
Continuous pilots are to be used to maintain flares. Two reliable methods of lighting
a flare pilot are the pressure ignitor and the electronic ignitor furnished by vendors
such as John Zink Company, National Airoil Burner Company, Flaregas and Kaldair.
In the past, operators used a Very pistol (which shoots balls of fire), Roman candles,
or a flaming arrow to ignite the flare pilots or burner.
The function of the ignitor is to ignite the pilots, which in turn light the flare burner.
Most flare tips have a minimum of three pilots equally spaced on the circumference
of the tip. These pilots are equipped with wind shields so that the most severe wind
cannot blow them out.
9.9.2
Pressure Ignitor
Each pilot has an ignitor pipe adjacent to it. A flame front is pushed through one line
from a remote ignitor to the top of the stack where the flame front is split to ignite all
the pilots.
The minimum pressure for fuel gas and air in the ignition chamber is 20 psi (1.4 bar,
1.4 kg/cm2). The minimum flows are 60 SCFH (1.7 Nm3/h) of fuel gas and 600
SCFH (17Nm3/h) of air. The flame front line size is usually 1 inch.
The fuel gas requirement for each pilot is 99,200 Btu/h (104,700 kJ/h, 25,000 kcal/h)
when the wind velocity is below 50 mph (80 km/h) and 19,800 Btu/h (209,300 kJ/h,
50,000 kcal/h) when the wind velocity is 50 mph (80 km/h) or more. The minimum
pressure is 30 psi (2.1 bar, 2.1 kg/cm2 ) downstream of the pilot gas pressure control
valve.
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A pressure ignitor can be located as far as 4,921 ft (1,500 m) from the pilots. A
vendor must be consulted to confirm this maximum distance. Typically, a pressure
ignitor is located as close to the flare stack as practical, but outside the safe radius
circle.
9.9.3
Electronic Ignitor
Electronic ignition systems have gained reliability in the last few years and should be
considered. They provide ignition of flare gas, pilot monitoring and automatic
reignition on loss of pilot.
The traditional method of lighting with a flame front ignitor requires several steps.
Electronic ignitors eliminate the need for flame front generation by introducing a
spark within the pilot nozzle. These systems are available either for manual or
automatic operation. In manual operation, ignition is initiated via a push-button after
the pilot flame is observed to be out. Automatic systems detect the presence of the
pilot flame via induced currents between the ignitor and ground. If the system
detects the pilot is lost, a spark is automatically released to reignite the pilot.
9.9.4
Atmospheric Ignitor
An atmospheric type ignitor may be used if plant compressed air is not available.
However, it is preferable to provide an air blower at the flare so that a pressured
ignitor can be used rather than an atmospheric ignitor.
The fuel gas pressure is regulated to about 5 psi (0.34 bar, 0.35 kg/cm2).
Atmospheric air for the ignition mixture is pulled into the igniting chamber by an
aspirator. The flame front line is usually a 2 inch or 3 inch pipe instead of a 1 inch
pipe as with the pressured ignitor. The operation of this ignitor is the same as with
the pressured ignitor except that the purging of the flame front line requires more
care. One disadvantage of this type of ignitor is that the aspirator requires careful
adjustment of the air vents and also must be kept very clean.
An atmospheric type ignitor can be located up to 330 ft (100 m) from the pilots. The
maximum distance must be confirmed by a vendor.
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9.9.5
Pilot Monitoring
An alarm system warns that a pilot at the burner is out. For pressure ignitors, this
consists of a chromel-alumel thermocouple installed at the base of each pilot and
connected to a temperature alarm switch. Loss of flame temperature then actuates
an alarm switch. Electronic ignitors can be installed with a self contained monitoring
system (see Section 9.9.3). Pilot monitoring should be installed for each flare.
9.10
REFERENCES
1.
Bleakley, W.B., “BP Flare System Gains Acceptance”, Petroleum Engineer
International, 25-28 (August 1978).
2.
Brzustowski, T.A., “A Model for Predicting the Shapes and Lengths of Turbulent
Diffusion Flames Over Elevated Industrial Flares”, paper presented at the 22nd
Canadian Chemical Engineering Conference, Toronto (1972).
3.
Brzustowski, T.A., “Flaring: The State of the Art”, A.I.Ch.E. 11th Annual Loss
Prevention Symposium, Houston (March 20-24, 1977).
4.
Brzustowski, T.A., and Sommer, E.C., Jr., “Predicting Radiant Heating from Flares”,
Proceedings Division of Refining, A.P.I., 53, 865-893 (1973).
5.
Brzustowski, T.A., Gollahalli, S.R., and Sullivan, H.F., “Characteristics of a Turbulent
Propane Diffusion Flame in a Cross-Wind”. 25th Canadian Ch.E. Conference,
Montreal, Canada, November 2-5, 1975, Rev. November 21, 1976, No. 75-CSME63, EIC Accession No. 1682 - (Abstracted on Transaction of the CSME, Volume 3,
No. 4, 1975).
6.
Cranfield, John, “BP’s New Indair Flare Burns Waste Gas Without Smoke”,
Petroleum and Petrochemical International, 86-89 (November, 1973).
7.
Ito, T. and Sawada, N., “Ground Flares Aid Safety”, Hydrocarbon Processing, 175190 (June, 1976).
CHAP9-R1.DOC
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8.
Kent, G.R., “Practical Design of Flare Stacks”, Hydrocarbon Processing, pp. 121-125
(August, 1964).
9.
Reed, R.D., Furnace Operations, Gulf Publishing, 12-31 (1973).
10.
Straitz, John F. III, “Make the Flare Protect the Environment”, Hydrocarbon
Processing, (October, 1977).
11.
Tan, S.H., “Flare System Design Simplified”, Hydrocarbon Processing, 172-176
(January, 1967).
12.
Vanderlinde, L.G., “Smokeless Flares”, Hydrocarbon Processing, 99-104 (October,
1974).
13.
Trade literature from the following flare design specialists:
•
John Zink Company
•
Kaldair Limited
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TABLE 9.1
COMPARISON OF FLARE TYPES
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TABLE 9.2A
LOWER LIMITS OF FLAMMABILITY
OF GASES AND VAPORS IN AIR
COMPOUND
SYMBOL
EMPIRICAL
FORMULA
LOWER LIMIT OF
FLAMMABILITY
(MOLE OR VOLUME %)
C1
C2
CH4
C2H6
5.00
3.22
Paraffin Hydrocarbons
Methane
Ethane
Propane
C3
C3H8
2.37
Butane
Isobutane
NC4
IC4
C4H10
C4H10
1.86
1.80
Pentane
Isopentane
C5
IC5
C5H12
C5H12
1.40
1.32
Hexane
Heptane
C6
C7
C6H14
C7H16
1.25
1.00
Octane
C8
C8H18
0.84
Nonane
Decane
C9
C10
C9H20
C10H22
0.74
0.67
Ethylene
C2
C2H4
3.05
Propylene
C3
C3H6
2.00
Butene-1
1-C4
C4H8
1.65 ∅
Butene-2
*2-C4
C4H8
1.75 ∅
C2
C2H2
2.50 ∅
Benzene
Toluene
BZ
TOL
C6H6
C7H8
1.41
1.27
O-Xylene
OXYL
C8H10
1.00 ∅
Olefins
Acetylenes
Acetylene
=
Aromatics
(*) Non-standard Symbols
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TABLE 9.2B
LOWER LIMITS OF FLAMMABILITY
OF GASES AND VAPORS IN AIR
COMPOUND
SYMBOL
EMPIRICAL
FORMULA
LOWER LIMIT OF
FLAMMABILITY
(MOLE OR VOLUME %)
H
H2
4.00
NH3
NH3
15.50
CO
CO
12.50
*C2O
C2H4O
3.00 ∅
*C3O
C3H6O
2.00 ∅
Carbon Disulfide
CS2
CS2
1.25 ∅
Hydrogen Sulfide
H2S
H2S
4.30
Methyl Chloride
*C1CL
CH3Cl
8.25 ∅
Vinyl Chloride
*VICL
C2H3Cl
4.00 ∅
*C2CL2
C2H4Cl2
6.20 ∅
Hydrogen
Hydrogen
Nitrogen Compounds
Ammonia
Oxides
Carbon Monoxide
Ethylene Oxide
Propylene Oxide
Sulfides
Chlorides
Ethylene Dichloride
(*) Non-Standard Symbols
Source:
Values for components marked (∅) are taken from “Low-Pressure Flare Processing for
Energy Conservation” by K.J. Hebert, McGill, Inc., Tulsa Oklahoma presented at the 1984
Spring National Meeting of the A.I.Ch.E. at Anaheim, California on May 21, 1984.
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TABLE 9.3
RECOMMENDED SURFACE
EMISSIVITY VALUES (εεs)*
Aluminum (oxidized)
0.25
Brass
0.06
Carbon Steel
0.9
Cast Iron
0.8
Monel
0.5
Stainless Steel
0.7
(*)
Emissivity values vary depending upon surface condition (oxidized,
polished, etc.) and/or surface coatings. The values given above are generally
high and therefore will result in high surface temperatures being predicted.
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TABLE 9.4
AIR REQUIRED FOR STOICHIOMETRIC
COMBUSTION OF GASES
Component
Formula
Air Requirements
6
Moles Air/Mole
kg Moles Air / 10 kcal
Component
Paraffins
Methane
CH4
9.5
3.16
Ethane
Propane
C2H6
C3H8
16.7
23.8
3.10
3.10
Butane
Pentane
C4H10
C5H12
31.0
38.1
3.10
3.10
Hexane
Heptane
C6H14
C7H16
45.2
52.4
3.08
3.08
Octane
C8H18
59.5
3.08
Nonane
Decane
C9H20
C10H22
66.7
73.8
3.08
3.08
Ethylene
Propene
C2H4
C3H6
14.3
21.4
2.87
2.95
Butane
C4H8
28.6
3.00
Pentene
C5H10
35.7
3.01
Acetylene
C2H2
11.9
2.52
Carbon Monoxide
Hydrogen Sulfide
CO
H2S
2.4
7.1
2.23
3.66
Ammonia
NH3
3.6
3.00
Hydrogen
Butadiene
H2
C4H6
2.4
26.2
2.62
2.78
Benezene
Toluene
C6H6
C7H8
35.7
42.9
3.02
3.05
C8H10
C2H4O
50.0
11.9
3.05
2.69
Olefins
Others
Xylene
Ethylene Oxide
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FIGURE 9.1
STACK & FLAME GEOMETRY
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FIGURE 9.2
XL VERSUS SL
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FIGURE 9.3
TEMPERATURE OF STEEL
VS TIME OF EXPOSURE
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FIGURE 9.4
TYPICAL ENCLOSED GROUNDFLARE CONFIGURATION
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FIGURE 9.5
TYPICAL BURN PIT
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FIGURE 9.6
TYPICAL AIR-ASSISTED FLARE SYSTEM
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FIGURE 9.7
STEAM/HYDROCARBON RATIO VERSUS
FLARE GAS MOLECULAR WEIGHT
FOR SMOKELESS FLARING
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FIGURE 9.8
CONVENTIONAL PIPE FLARE
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FIGURE 9.9
CONVENTIONAL FLARE
WITH STEAM WATER SPRAY
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FIGURE 9.10
HIGH VELOCITY TIP
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FIGURE 9.11
AIR ASSISTED FLARE TIP
(TOP VIEW)
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FIGURE 9.12
COANDA NOZZLE (INTERNAL)
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FIGURE 9.13
COANDA FLARE (EXTERNAL)
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FIGURE 9.14
OFFSHORE FLARE SUPPORT TYPES
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APPENDIX
A
APPENDIX A
PAGE
1
NOMENCLATURE
DATE
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APPENDIX A
NOMENCLATURE
Accumulation
The pressure increase over the maximum allowable working pressure of the vessel during discharge
through the pressure relief device, expressed in pressure units or as a percent. Maximum allowable
accumulation is established by applicable codes for operating and fire contingencies.
Actual Discharge Area
The measured minimum net area that determines the flow through a valve.
Back Pressure
The pressure that exists at the outlet of a pressure relief device as a result of the pressure in the
discharge system. It is the sum of the superimposed and built-up back pressure.
Balanced Pressure Relief Valve
A spring-loaded pressure relief valve that incorporated a means for minimizing the effect of back
pressure on the performance characteristics.
Blowdown
The difference between the set pressure and the closing pressure of a pressure relief valve,
expressed as a percentage of the set pressure or in pressure units.
Built-up Back Pressure
The increase in pressure in the discharge header that develops as a result of flow after the pressure
relief device opens.
Closing Pressure
The value of decreasing inlet static pressure at which the valve disk reestablishes contact with the
seat or at which lift becomes zero.
Cold Differential Test Pressure
The pressure at which the pressure relief valve is adjusted to open on the test stand. The cold
differential test pressure includes corrections for the service conditions of back pressure or
temperature or both.
Conventional Pressure Relief Valve
A spring-loaded pressure relief valve whose performance characteristics are directly affected by
changes in back pressure on the valve.
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APPENDIX A
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NOMENCLATURE
DATE
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Curtain Area
The area of the cylindrical or conical discharge opening between the seating surfaces above the
nozzle seat created by the lift of the disk.
Design Gauge Pressure
The most severe conditions of coincident temperature and pressure expected during operation. This
pressure may be used in place of the maximum allowable working pressure in all cases where the
MAWP has not been established. The design pressure is equal to or less than the MAWP.
Effective Discharge Area
The nominal or computed area of a pressure relief valve used in recognized flow formulas to
determine the size of the valve. It will be less than the actual discharge area.
Huddling Chamber
An annular pressure chamber in a pressure relief valve located beyond the seat for the purpose of
generating a rapid opening.
Inlet Size
The nominal pipe size (NPS) of the valve at the inlet connection, unless otherwise designated.
Leak-Test Pressure
The specified inlet static pressure at which a seat leak test is performed.
Lift
The actual travel of the disk away from the closed position when a valve is relieving.
Maximum Allowable Working Pressure (MAWP)
The maximum gauge pressure permissible at the top of a completed vessel in its operating position
for a designated temperature. The pressure is based on calculations for each element in a vessel
using nominal thicknesses, exclusive of additional metal thicknesses allowed for corrosion and
loadings other than pressure. The maximum allowable working pressure is the basis for the pressure
setting of the pressure relief devices that protect the vessel.
Maximum Operating Pressure
The maximum pressure expected during system operation.
Nozzle Area
The cross-sectional flow area of a nozzle at the minimum nozzle diameter.
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Opening Pressure
The value of increasing inlet static pressure at which there is a measurable lift of the disk or at which
discharge of the fluid becomes continuous.
Outlet Size
The nominal pipe size (NPS) of the valve at the discharge connection, unless otherwise designated.
Overpressure
The pressure increase over the set pressure of the relieving device, expressed in pressure units or as
a percent. It is the same as accumulation when the relieving device is set at the maximum allowable
working pressure of the vessel.
Pilot-Operated Pressure Relief Valve
A pressure relief valve in which the main valve is combined with and controlled by an auxiliary
pressure relief valve.
Pressure Relief Device
A device actuated by inlet static pressure and designed to open during an emergency or abnormal
conditions to prevent a rise of internal fluid pressure in excess of a specified value. The device may
also be designed to prevent excessive internal vacuum. The device may be a pressure relief valve, a
nonreclosing pressure relief device, or a vacuum relief valve.
Rated Relieving Capacity
That portion of the measured relieving capacity permitted by the applicable code or regulation to be
used as a basis for the application of a pressure relief device.
Relief Valve
A spring loaded pressure relief valve actuated by the static pressure upstream of the valve. The valve
opens normally in proportion to the pressure increase over the opening pressure. A relief valve is
used primarily with incompressible fluids.
Relieving Conditions
Used to indicate the inlet pressure and temperature of a pressure relief device at a specified
overpressure. The relieving pressure is equal to the valve set pressure (or rupture disk burst
pressure) plus the overpressure. (The temperature of the flowing fluid at relieving conditions may be
higher or lower than the operating temperature.)
Rupture Disk Device
A nonreclosing differential pressure relief device actuated by inlet static pressure and designed to
function by bursting the pressure containing rupture disk. A rupture disk device included a rupture
disk and a rupture disk holder.
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Safety Valve
A spring-loaded pressure relief valve actuated by the static pressure upstream of the valve and
characterized by rapid opening or pop action. A safety valve is normally used with compressible fluids.
Safety Relief Valve
A spring-loaded pressure relief valve that may be used as either a safety or relief valve depending on
the application.
Set Pressure
The inlet gauge pressure at which the pressure relief valve is set to open under service conditions.
Simmer
The audible or visible escape of compressible fluid between the seat and disk at an inlet static
pressure below the set pressure and at no measurable capacity.
Spring Loaded Pressure Relief Valve
A pressure relief device designed to automatically reclose and prevent the further flow of fluid.
Stamped Capacity
The rated relieving capacity that appears on the device nameplate. The stamped capacity is based
on the set pressure or burst pressure plus the allowable overpressure for compressible fluids and the
differential pressure for incompressible fluids.
Superimposed Back Pressure
The static pressure that exists at the outlet of a pressure relief device prior to lifting as a result of the
pressure in the discharge system.
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APPENDIX B
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TOWER RELIEF LOAD
DATE
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APPENDIX B-1
TOWER RELIEF LOAD
LIMITATIONS
There are some cases for which this calculation method may not be directly applicable. Please
read special considerations in Section 4.3.5.5 for the following tower systems:
•
Steam stripped fractionators
•
Non-reboiled fractionators with heavy bottoms such as crude/vacuum towers
•
Strippers for absorption systems
SPREADSHEET OVERVIEW
The “Tower Relief Load Calculations” Excel workbook was created to facilitate relief load
calculations. Calculations are based on the criteria defined in Section 4.3. The equations are
included before the spreadsheets under Backup Calculations. The workbook consists of 12
working spreadsheets interfacing with users, and 9 hidden spreadsheets for storing supporting
macros that make the workbook user-friendly. The first 12 spreadsheets consist of the following
spreadsheets:
Normal Case (Sheet 1) – Establishes normal heat and material balance
Relief Cases (Sheets 2-9) – Calculates relief loads for various failure scenarios
Feed/Bottoms Heat Exchanger Rating (Sheet 10) – Finds new feed temperature & enthalpy
Reboiler Thermal Rating (Sheet 11) – Determines pinch feasibility
Accumulator Surge Time (Sheet 12) – Determines time before accumulator floods
A copy of these Excel spreadsheets can be found under Flare System Calculation Spreadsheets in
the Fluor Daniel Intranet/Functions/Engineering/Disciplines/ Process/Aliso Viejo/Process
Knowledge/ Process Design Manuals.
APXB1-R1.DOC
APPENDIX
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TOWER RELIEF LOAD
DATE
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Before using this spreadsheet, the user must establish a Heat and Material Balance Envelope.
This envelope should be drawn so that all streams leave as saturated liquid or saturated vapor.
This envelope should include tower overhead condenser, accumulator and reflux system, tower
bottoms reboiler, exchanger or fired heater, pumparound cooling system, preheater (if a fired
heater is located at the tower inlet) and side-draw strippers and surge pots. Do not include
feed/bottoms exchangers or product coolers in the envelope and assume the overhead product
leaves from the reflux drum as a saturated liquid (i.e. disregard the pump head).
The function of the worksheets will be described in this overview.
Normal Case - Tower Relief Load Calculations (Sheet 1)
This spreadsheet reports general information about the PSV and the tower design basis as
well as establishes a normal heat and material balance. As with all spreadsheets in the
workbook, the user should only input the items in blue. The following General Information
must be input: Plant, Unit, Equipment Number, Equipment Serviced, P&ID, PSV Tag, Set
Pressure, and Over Pressure (Allowable Over Pressure = 10% for single valve, 16% for
multiple valves). The Unit and Tower Design Basis section is only for reference. Data in
this section is not used for calculations. The Menu button at the top of the sheet can be
used to do the following: Set Units (English or Metric), Set Unit Design Basis, Set Tower
Design Basis, View Exchanger Rating, View Reboiler Thermal Rating, View Accumulation
Surge Time, and Print Report. Input a Pro-ration factor applied to simulation. Using the
PFD markup with the heat and material balance envelope, list all the streams and the
corresponding flow rates, phases, °API, MW, temperatures, pressures, enthalpies, and
duties. Enter a BASIS FACTOR of 1 for each stream and verify that both material and
energy are balanced. The Message cell will display a warning message to check the
balance if the IN and OUT TOTALS do not match within 1%. Finally, document the BASIS
FACTORS and add any appropriate NOTES in the appropriate sections. General
information and data in this spreadsheet will be automatically transferred to the following 8
relief case spreadsheets.
APXB1-R1.DOC
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DATE
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Relief Cases - Tower Relief Load Calculations (Sheets 2 - 9)
A total of 8 spreadsheets are included for various relief cases. If a tower has more than 8
relief cases, the user can copy these spreadsheets for the extra cases. The first relief case
spreadsheet requires the user to input the relief conditions and properties for the streams.
These conditions and properties will be automatically transferred to the other 7 relief cases.
With these 7 relief cases, the user is only required to enter the case description, assign
basis factors, and document those factors. The properties of the OVHD and BOTTOMS
PRODUCTS can be obtained from a simulation. In the event that a feed/bottoms
exchanger exists, a new temperature and corresponding enthalpy is determined using
Feed/Bottoms Heat Exchanger Rating (Sheet 10). The BASIS FACTORS for each stream
are then adjusted to account for the failure scenario. A material balance must always be
maintained. The Accumulation liquid and vapor enthalpies (obtained from a dew point flash
of overhead liquid product or reflux) must also be entered for the sheet to calculate a latent
heat of vaporization. The MW, temperature, heat capacity ratio k=Cp/Cv, compressibility
factor, and viscosity of the Accumulation vapor must also be entered by the user. The
sheet then subtracts the Total Energy Out from the Total Energy In and divides the result
by the Latent Heat of Vaporization to determine the ACCUMULATION. The flow rates of
the OVHD PRODUCT, OFF GAS and STEAM are also listed. The spreadsheet then adds
ACCUMULATION and OVHD PROD. If this sum is less than or equal to 0, the RELIEF
LOAD is the sum of the OFF GAS and STEAM. If the sum is greater than 0, the RELIEF
LOAD is the sum of the ACCUMULATION, OVHD PROD, OFF GAS, and STEAM. The
user is responsible for determining the appropriate BASIS FACTORS and documenting
them in the BASIS section. When determining the BASIS FACTOR for the REBOILER, it is
important to determine the feasibility of a reboiler pinch using the Reboiler Thermal Rating
(Sheet 11) spreadsheet. When determining the BASIS FACTOR(s) for OVHD
CONDENSER(S), it is important to determine if the accumulator floods. If the accumulator
floods in less than 10 minutes, it is assumed that the condenser will also flood and the duty
will be reduced to 0. The Accumulator Surge Time (Sheet 12) spreadsheet calculates
approximate surge time based on the vessel dimensions during relief scenarios.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
4
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Feed/Bottoms Heat Exchanger Rating (Sheet 10)
This spreadsheet calculates a new feed temperature and enthalpy during a relief scenario
where a feed/bottoms exchanger is present. The calculation assumes that the UA value is
constant and heat transfer rate is constant as of normal operation and temperature change
is linear (phase change does not occur). The sheet requires the user to enter the names of
the Tower Feed and Bottoms, the Normal Feed inlet temperature and enthalpy, the Normal
Bottoms outlet temperature, and the Normal Duty. The calculation sequence and equations
are shown in the Calculation Backup section included after the spreadsheets. The
spreadsheet instructs the user how to use the spreadsheet. Once the program has
converged, go to the relief case where Feed/Bottoms Heat Exchanger Rating applies and
enter the new feed temperature and enthalpy.
Reboiler Thermal Rating (Sheet 11)
This spreadsheet determines approximately, but conservatively the feasibility of taking
reboiler pinch credit. If the clean surface heat transfer with reduced LMTD is less than the
normal duty, consider a reboiler pinch. Reboiler pinch calculations are based on the
following assumptions:
(1) clean reboiler surface (clean U)
(2) bottoms enthalpy rise is proportional to bottoms temperature rise
(3) LMTD correction factor (f) is 1
The sheet requires the user to enter the tower bottoms tin, the reboiler Tin and Tout, the clean
heat transfer coefficient, and the surface area of the reboiler. All these values are at the
normal case conditions. If the reboiler uses steam, “Yes” should be entered in the
appropriate cell. The calculation sequence and equations are shown in the Calculation
Backup section included after the spreadsheets. The spreadsheet instructs the user how to
use the spreadsheet. Once the program has converged, go to the relief case where
reboiler thermal rating applies and enter the Basis Factor.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
5
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Accumulator Surge Time (Sheet 12)
This spreadsheet determines the surge time before the accumulator and condenser flood.
The sheet requires the user to enter the equipment name and number, P&ID number,
vessel orientation (1= HORIZONTAL, 2=VERTICAL), diameter, length T/T, liquid level,
basis notes and factors, density at relief conditions, and the status and flow rate of the
reflux and overhead product flows. The calculation sequence and equations are shown in
the Calculation Backup section after the spreadsheets. Volume calculations assume
elliptical 2:1 heads. If the accumulator surge time is less than 10 minutes, adjust the
corresponding relief case basis factors of the condensers to 0 because the condenser will
flood.
SAMPLE CALCULATION
PROBLEM DEFINITION
This example illustrates calculation methods presented in Section 4.3. These calculations
can be easily performed using the attached spreadsheets. A Debutanizer Column is fed
from a Stripper Column. HCO Pumparound from the main fractionator column is used to
reboil the column and an air cooler, followed by a trim cooler, is used to condense the
overhead vapor. The reflux/product pump motor service is split between two 34.5 KV
buses, with one pump motor in each service being supplied from each Bus. Air cooler fan
motors are also equally split between the two buses. There are 4 different partial power
failure scenarios. Calculate the relief loads associated with total and partial power failures,
cooling water failure, and reflux failure.
REQUIRED DATA
All streams in and out of the envelope require molecular weight and enthalpy at the normal
and relieving conditions. All the data required for the calculations are summarized in Table
B1-1.
•
Data at normal operation is obtained from the design heat and material balance.
•
Exchanger duties are based on the normal case heat balance.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
6
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
•
Power failure modes are derived from electrical one-line diagrams.
•
Accumulator dimensions are obtained from vessel drawings.
•
Relief conditions are determined at the Relief Pressure: Relief Pressure = Set Pressure
(185 PSIG) + Allowable Overpressure (16%) + Atmospheric Pressure (14.70 PSIA) =
229.3 PSIA.
•
Data at relief conditions is obtained from stream simulation at the relief conditions as
follows:
•
Bottoms Product: Obtain temperature and enthalpy from a bubble point flash
calculation at the tower bottom pressure of 155 psig. If there are side products,
calculate the required data in the same way. Note: This method can not be
used for a tower using significant stripping steam and non-condensable gas
such as an absorber. A better result can be obtained by using the Cox Chart. In
this case, estimate the relief temperature by extrapolating to relief pressure and
then calculate enthalpy at the relief temperature by a stream simulation.
•
Overhead Liquid Product: Assume that this stream leaves the envelope as a
saturated vapor for calculation purposes. This vapor is added to relief load at
the end of the load calculation. Obtain the data from a dew point flash
calculation at relief pressure of 150 psig.
•
Accumulation: Accumulation is a hypothetical stream that is assumed to be
generated from the top tray by an energy imbalance during a relief scenario.
Use the dew point conditions from the simulation performed for the Overhead
Liquid Product. The latent heat of the accumulation is calculated from the dew
point flash calculation (vapor enthalpy – liquid enthalpy).
•
Overhead Vapor Product: Not applicable in this case. If there is a normal vapor
product, obtain data from a dew point flash calculation of this stream.
•
Feed: The enthalpy of the feed stream is corrected for the potential duty
increase at the Feed/Bottoms exchanger due to the bottom temperature
increase during relief scenario. Use the attached heat exchanger rating
spreadsheet to determine corrected feed temperature and enthalpy.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
7
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
TOWER RELIEF LOAD CALCULATIONS
TOTAL POWER FAILURE
The Basis Factors (BF) are determined as follows:
•
Feed coming from Stripper C-42515 continues to flow at normal rates to the
Debutanizer until the Stripper is empty. FEED BF = 1.00.
•
Reboiler has no heat input due to loss of power for HCO Pumparound. REBOILER BF
= 0.00.
•
Overhead product must continue flowing to maintain material balance. OVHD
PRODUCT BF = 1.00.
•
Bottoms product exits at normal rate, but at relieving conditions. This will cause the
heat transfer of the feed/bottoms exchanger to increase and, thus, increase the feed
temperature and enthalpy. BOTTOMS PRODUCT BF = 1.00.
•
With the condenser fans stopped, the condenser duty is reduced to 25% due to natural
convection. CONDENSER BF = 0.25.
•
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water
pump. TRIM CONDENSER BF = 0.50.
•
At this overall duty and loss of reflux pump, the surge time is less than 10 minutes, and,
therefore, no credit is taken for condensing. CONDENSER BF = 0.00 and TRIM
CONDENSER BF = 0.00
These inputs result in a calculated loss of heat from the column, which indicates there is no
relief load.
PARTIAL POWER FAILURE
Feed Stops and Reflux Pump Stops (PPF-1):
The Basis Factors (BF) are determined as follows:
•
Feed from Stripper C-42515 stops due to stripper feed stop. FEED BF = 0.00.
•
Reboiler continues. REBOILER BF = 1.00.
•
Overhead and bottoms products must stop to maintain material balance. OVHD
PRODUCT BF = 0.00 and BOTTOMS PRODUCT BF = 0.00.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
8
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
•
Loss of two of the four condenser fans results in a condensing duty of 62.5% overall.
(100% for the half with the fans running and 25% for the half with natural convection
only). CONDENSER BF = 0.63
•
The trim condenser duty is reduced by 50% due to loss of motor driven cooling water
pumps. TRIM CONDENSER BF = 0.50.
•
However, the Reflux/Product pump is shutdown so the condensers will flood and no
credit can be taken for condenser duty. CONDENSER BF = 0.00 and TRIM
CONDENSER BF = 0.00.
These inputs result in a relief load of 684,032 lb/hr.
Feed Continues and Reflux Pump Stops (PPF-2):
The Basis Factors (BF) are determined as follows:
•
Feed coming from Stripper C-42515 continues to flow at normal rates to the
Debutanizer. FEED BF = 1.00.
•
Reboiler continues. REBOILER BF = 1.00.
•
Overhead product must continue flowing to maintain material balance. OVHD
PRODUCT BF = 1.00.
•
Bottoms product exits at normal rate, but at relieving conditions. This will cause the
heat transfer of the feed/bottoms exchanger to increase and, thus, increase the feed
temperature and enthalpy. BOTTOMS PRODUCT BF = 1.00.
•
Loss of two of the four condenser fans results in a condensing duty of 62.5% overall.
(100% for the half with the fans running and 25% for the half with natural convection
only). CONDENSER BF = 0.63
•
The trim condenser duty is reduced by 50% due to loss of motor driven cooling water
pumps. TRIM CONDENSER BF = 0.50.
•
However, the Reflux/Product pump is shutdown so the condensers will flood and no
credit can be taken for condenser duty. CONDENSER BF = 0.00 and TRIM
CONDENSER BF = 0.00.
These inputs result in a relief load of 571,463 lb/hr.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
9
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Feed Continues and Reflux Pump Continues (PPF-3):
The Basis Factors (BF) are determined as follows:
•
Feed coming from Stripper C-42515 continues to flow at normal rates to the
Debutanizer. FEED BF = 1.00.
•
Reboiler continues. REBOILER BF = 1.00.
•
Overhead product must continue flowing to maintain material balance. OVHD
PRODUCT BF = 1.00.
•
Bottoms product exits at normal rate, but at relieving conditions. This will cause the
heat transfer of the feed/bottoms exchanger to increase and, thus, increase the feed
temperature and enthalpy. BOTTOMS PRODUCT BF = 1.00.
•
Loss of two of the four condenser fans results in a condensing duty of 62.5% overall.
(100% for the half with the fans running and 25% for the half with natural convection
only). CONDENSER BF = 0.63
•
The trim condenser duty is reduced by 50% due to loss of motor driven cooling water
pumps. TRIM CONDENSER BF = 0.50.
These inputs result in a relief load of 163,617 lb/hr.
Feed Stops and Reflux Pump Continues (PPF-4):
The Basis Factors (BF) are determined as follows:
•
Feed stops. FEED BF = 0.00.
•
Reboiler has no heat input due shutdown of HCO Pumparound. REBOILER BF = 0.00.
•
Overhead and bottoms product must stop to maintain material balance. OVHD
PRODUCT BF = 0.00 and BOTTOMS PRODUCT BF = 0.00.
•
Loss of two of the four condenser fans results in a condensing duty of 62.5% overall.
(100% for the half with the fans running and 25% for the half with natural convection
only). CONDENSER BF = 0.63
•
The trim condenser duty is reduced by 50% due to loss of motor driven cooling water
pumps. TRIM CONDENSER BF = 0.50.
These inputs result in a relief load of 0 lb/hr.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
10
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
COOLING WATER FAILURE
The Basis Factors (BF) are determined as follows:
•
Feed coming from Stripper C-42515 continues to flow at normal rates to the
Debutanizer. FEED BF = 1.00.
•
Reboiler continues. REBOILER BF = 1.00.
•
Overhead product must continue flowing to maintain material balance. OVHD
PRODUCT BF = 1.00.
•
Bottoms product exits at normal rate, but at relieving conditions. This will cause the
heat transfer of the feed/bottoms exchanger to increase and, thus, increase the feed
temperature and enthalpy. BOTTOMS PRODUCT BF = 1.00.
•
Condenser fans are operational. CONDENSER BF = 1.00
•
Loss of cooling water supply will cause the trim cooler to shutdown. TRIM
CONDENSER BF = 0.00
These inputs result in a relief load of 72,562 lb/hr.
REFLUX FAILURE
The Basis Factors (BF) are determined as follows:
•
Feed coming from Stripper C-42515 continues to flow at normal rates to the
Debutanizer. FEED BF = 1.00.
•
Reboiler continues. REBOILER BF = 1.00.
•
Overhead product must continue flowing to maintain material balance. OVHD
PRODUCT BF = 1.00.
•
Bottoms product exits at normal rate, but at relieving conditions. This will cause the
heat transfer of the feed/bottoms exchanger to increase and, thus, increase the feed
temperature and enthalpy. BOTTOMS PRODUCT BF = 1.00.
•
Loss of Reflux/Product pumps will cause the condensers to flood. CONDENSER BF =
0.00 and TRIM CONDENSER BF = 0.00.
These inputs result in a relief load of 571,463 lb/hr.
REBOILER PINCH CALCULATION
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
11
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Using the Reboiler Thermal Rating spreadsheet, the reboiler pinch calculation indicates that
there is no feasibility of reboiler pinch.
CONCLUSIONS
Based on the calculations shown in the attached spreadsheets, the total power failure load
is zero and the maximum relief load resulting from partial power failure load is about
684,032 lb/hr. Cooling water failure results in a relief load of 72,562 lb/hr. Reflux failure
results in a relief load of 571,463 lb/hr.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
12
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
TABLE B1-1
INPUT DATA
(Note: This input data sheet is not a required document. The required data may be directly entered
into the spreadsheets from the simulation and other sources.)
Operating Column Pressure: Top – 150 PSIG, Bottom – 155 PSIG
Operating Accumulator Pressure: 140 PSIG
Set Pressure: 185 PSIG
Relief Pressure: 229.3 PSIA
Stream
Flowrate
MW
LB/HR
Temp.
Press.
Enthalpy
°F
PSIA
BTU/LB
Normal Case
Feed
728,303
86.6
288.1
167.7
158.26
Overhead Product
144,908
50.5
111.0
154.7
45.70
Bottoms Product
583,395
105.3
382.0
169.7
186.20
Feed (*)
728,303
86.6
301.1
232.2
177.32
Overhead Product
144,908
50.5
178.4
229.3
207.09
Bottoms Product
583,395
105.3
435.7
234.3
222.11
Top Tray Liquid
53.8
178.4
229.3
88.85
Top Tray Vapor
50.5
178.4
229.3
207.09
Relief Case
(*)
Feed temperature is corrected or the increased heat transfer of the feed/bottom exchanger
during relief.
Normal Reboiler Duty
80.88 MMBtu/hr
Normal Condenser Duty
58.99 MMBtu/hr
Normal Trim Condenser Duty
22.12 MMBtu/hr
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
13
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Top Tray
Liquid
k = Cp/Cv
1.2735
Z
0.7724
Viscosity, cP
0.0105
Reboiler
Area
8655.2 ft2
U (clean)
109.4 Btu/h ft2 °F
Normal Process-Side Temp.
In - 356 °F, Out - 382 °F
Normal HCO Pumparound Temp.
In - 580 °F, Out - 478 °F
Service
Total Power Failure
Partial Power Failure
1
2
3
4
Y
N
Y
N
Y
Reflux Pump
Off
On
Off
Off
On
Condenser Fans
Off
½
½
½
½
CW Pumps
½
½
½
½
½
HCO Pumparound
Off
Off
On
On
On
Feed Continues (Y/N)
Accumulator
Diameter
9.5 ft
Length, T/T
24.0 ft
NLL
5.25 ft
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
14
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
FIGURE B1.1
PROCESS SKETCH
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
15
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
BACKUP CALCULATIONS
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
16
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Feed/Bottoms Heat Exchanger Rating
TOWER
T out
t in
tout
HEAT
EXCHANGERS
T in
In order to use this spreadsheet, all bold values or blue-color values must be entered.
These values include: name of tower feed and bottoms streams, normal feed temperature
in (tin) and corresponding enthalpy, normal bottoms temperature out (Tout), and normal
exchanger duty (QN). The spreadsheet automatically downloads the normal tower feed
temperature out (tout) and corresponding enthalpy and flow rate, normal bottoms
temperature in (Tin) and flow rate, and calculates normal LMTD. The spreadsheets also
downloads the relief flow rates for both feed and bottoms and the bottoms relief
temperature in (Tin), and assumes that the relief feed temperature in and enthalpy are the
same as the normal case. The user must then enter a QA/QN (guess) of 1.25. If “REITERATE” appears, click on the re-iterate button or enter QA/QN (next guess) value for
QA/QN (guess) until the spreadsheet converges on a solution. The steps the spreadsheet
follows are shown below.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
17
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Useful Equations
Normal QN = U A ∆tN
Assume U A = constant
QN = W B CPB (Ti – To)
W B CPB = QN / ∆TN
QN = W F CPF (to – ti)
W F CPF = QN / ∆tN
LMTD = ((Ti – to) – (To – ti)) / ln((Ti – to) / (To – ti))
Calculation Sequence
Assume:
QA = 1.25 QN
Calculate:
To = Ti – QA/(W BCPB) = Ti – 1.25 QN / (QN / ∆TN) = Ti – 1.25 ∆TN
Calculate:
to = ti + QA/(W FCPF) = ti + 1.25 QN / (QN / ∆tN) = ti + 1.25 ∆tN
Calculate:
(LMTD)R = ((Ti – to) – (To – ti)) / ln((Ti – to) / (To – ti))
Using calculated To and to.
Calculate:
QR = QN (LMTD)R/(LMTD)N
IF |(QA / QR – 1)| ≤ 0.05, stop iteration and report QR.
IF QR ≥ 1.5 QN, report QR = 1.5 QN
If deviation is larger than 5%, repeat calculation by increasing and decreasing (QA).
The spreadsheet calculates the feed enthalpy as follows: CPFOUT = CPFIN + QR / W F
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
18
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Reboiler Thermal Rating
“Pinch Calculation”
TOWER
tout
REBOILER
T out
T in
tin
BOTTOMS PRODUCT
In order to use this spreadsheet, all bold values or blue-color values must be entered.
These values include: normal tower bottoms temperature in (tinN), reboiler temperature in
and out (TinN and ToutN), clean heat transfer coefficient (Uc), and heat exchanger surface
area (A). The user should indicate if the reboiler uses steam. The spreadsheet
automatically downloads the normal tower bottoms temperature out (toutN), the normal
reboiler duty (QN), and the relief tower bottoms temperature out (toutR). The reboiler relief
TinR, Uc, and A are assumed to be the same as the normal case. The spreadsheet checks
a zero reboiler case and a full normal reboiler duty case in order to determine whether
pinch calculations are required for the relief conditions (Equations 1 and 2).
If
Uc A (TinR – toutR) / QN ≤ 0, then zero reboiler duty
APXB1-R1.DOC
(Equation 1)
APPENDIX
B-1
APPENDIX B
PAGE
19
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
If
Uc A (ToutN – toutR) / QN ≥ 1, then full normal reboiler duty
(Equation 2)
Where:
Uc
=
Clean heat transfer coefficient
A
=
Heat exchanger surface area
TinR
=
Heat medium inlet temperature during relief.
Steam: saturated steam temperature at supply pressure
Heat Medium Liquid: heat medium supply temperature
toutR
=
Bottoms temperature during relief
QN
=
Normal reboiler duty
ToutN =
Normal heat medium outlet temperature
Otherwise, a reboiler pinch calculation is performed as shown below.
Assumptions:
1) Reboiler surface is clean, use clean U
2) Bottoms enthalpy rise is proportional to bottoms temperature rise
Steam Reboiler (commonly used for thermosiphon or kettle reboilers)
Useful Equations
QR
=
Uc A ∆tm
∆tB
=
to – ti
ti
=
to - ∆tB = tB - ∆tB
to
=
tB
∆tm
=
f (MTD)
MTD =
(∆ti + ∆to) / 2
=
((Ts – ti) + (Ts – to)) / 2
=
(Ts-tB+∆tB + Ts – tB)/2
=
(Ts – tB) + ½ ∆tB
APXB1-R1.DOC
Assume f = 1
Steam side Ti = To = Ts
APPENDIX
B-1
APPENDIX B
PAGE
20
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
QR1
=
Uc A (Ts – tB + ½ ∆tB)
(Equation 3)
QR2
=
QN (∆tB/∆tBN)
(Equation 4)
Where:
QN
=
Normal Reboiler duty from simulation
QR
=
Relief case reboiler duty
∆tBN
=
Normal temperature rise (Bottom two tray temperature difference from
simulation)
Ts
=
Saturated steam temperature at supply pressure
tB
=
Tower bottoms temperature (from relief load calculation sheet)
Uc
=
Clean U from data sheet
A
=
Surface area from data sheet
Calculation Sequence
First input ∆tB = ∆tBN in Equation 3.
Perform iteration by inputting the calculated QR1 into the following equation
∆tB
=
∆tBN (QR1 / QN)
Continue iterating until
(1 – QR1/QR2) ≤ ±0.05
Heat Medium Reboiler
Useful Equations
QR
=
Uc A ∆tm
∆tm
=
f (LMTD)
(Equation 5)
Assume f = 1
LMTD =
((To – ti) – (Ti – to)) / ln((To – ti) / (Ti – to))
Ti
=
heat medium supply temperature
to
=
tB = bottoms temperature at relief
QA/QN =
∆TR / ∆TN = (Ti – To ) / ∆TN
To
Ti – (QA/QN) ∆TN
=
APXB1-R1.DOC
(Equation 6)
(Equation 7)
APPENDIX
B-1
APPENDIX B
PAGE
21
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
QA/QN =
∆tB / ∆tN = (to – ti) / ∆tN
ti
to – (QA/QN) ∆tN
=
(Equation 8)
Calculation Sequence
Assume QA = ½ QN initially in Equation 7 and 8. Calculate (LMTD) in Equation 6 and
calculate QR from Equation 5. Compare this QR with the assumed one (QA).
If
|(QA/QR – 1)| ≤ 0.05,
stop iteration and report calculated QR or QR/QN
If
|(QA/QR – 1)| ≥ 0.05,
increase or decrease duty assumption (QA) by 10%.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
22
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Accumulator Surge Time
During a failure scenario, the liquid level in the overhead accumulator may rise and
flood the overhead condenser, reducing its duty to 0. The Accumulator Surge Time
spreadsheet calculates the amount of time available before the accumulator and
condenser flood. If the Accumulator Surge Time is less than 10 minutes, it is assumed
that the accumulator and condenser will flood before an operator can correct the situation.
The spreadsheet will automatically list the Relief Case Description and perform
calculations based on the values input by the user (blue text). The user is required to input
the Equipment Number, Equipment name, P&ID Number, Vessel Orientation
(1=Horizontal, 2=Vertical), accumulator diameter, length (tangent to tangent), liquid level,
basis factors, reflux and overhead product rates, and the accumulation density. The
condenser and trim condenser duties and latent heat of accumulation are automatically
obtained from the normal and total power failure cases. The Volumetric Flow Rate,
Condensables is obtained by dividing the Total Condenser Duty by the Latent Heat and
then by the Accumulation Density. The Total Volumetric Flow Rate is obtained by
subtracting the Reflux and/or the Overhead Product if those flows are indicated to
continue. The Accumulator Total and Partial Volumes are determined by the equations
below.
Horizontal Cylindrical Tanks
Total volume
=
volume in 2 heads + volume in cylinder
=
1/6 π K1 D3 + ¼ π D2 L
K1
=
2 b/D
Ze
=
H1/D
Zc
=
H1/D
=
1/6 π K1 D3 [f(Ze)] + ¼ π D2 L [f(Zc)]
Partial Volume
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
23
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
F(Zc) =
Horizontal cylinder coefficient
F(Ze) =
Ellipsoidal coefficient
For elliptical 2:1 heads, b = ¼ D, K1 = ½
Vertical Cylindrical Tanks
Total Volume
Partial Volume
=
volume in heads + volume in cylinder
=
1/6 π K1 D3 + ¼ D2 L
=
Assume that if vessel contains liquid, it will fill one elliptical
head. If vessel liquid level is same as tangent to tangent
length, assume that both heads are also full.
If H = L
=
Total Volume
If H < L
=
1/6 π K1 D3 (0.5) + ¼ π D2 H
=
2 b/D
K1
For elliptical 2:1 heads, b = ¼ D, K1 = ½
The Accumulator Surge Volume is the difference between the Total Volume and the Partial
Volume. The Accumulator Surge Time is determined by dividing the Accumulator Surge
Volume by the Total Volumetric Flow Rate. If the surge time is close to 10 minutes, a
more detailed evaluation is recommended.
APXB1-R1.DOC
APPENDIX
B-1
APPENDIX B
PAGE
24
TOWER RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Mixture Properties
Once a relief load is calculated, the load and properties are reported. In the case of
mixtures, use the following formulas to determine properties.
MW = Σ W i / Σ (W/MW)i
T = Σ W iT i / Σ W i
µ = Σ xi µI (MW)i
0.5
/ Σ xi (MW)i
0.5
i
=
ith component
W
=
Mass flow rate
X
=
Mass fraction
T
=
Temperature
MW
=
Molecular weight
µ
=
Viscosity
th
PG 5-16, GPSA, 10 Edition, 1994
Z = Σ yiZi
yi
Z
=
Mole fraction of component in vapor
=
Compressibility of component
th
Perry’s, 6 Edition, 1984, pg 3-280
K
Σyiki
=
K
=
th
Heat capacity ratio of component i
Perry’s, 6 Edition, 1984, pg 3-275
APXB1-R1.DOC
Contract :
4499920
Revision :
Date :
By:
Chk:
A
26-Jun-2000
JLH
DK
Appv'd:
DDC
Page 1 of 12
TOWER RELIEF LOAD CALCULATIONS
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
GENERAL
TOWER DESIGN BASIS
UNIT FEED RATE
TOWER FEED RATE
70,000 BPD @ 60°F
70,000 BPD @ 60°F
464 KL/H @ 60°F
464 KL/H @ 60°F
7,952 T/D
7,952 T/D
66.5 °API
66.5 °API
Relief Pressure = Set Pressure + Allowable Overpressure + Atmospheric Pressure (14.696 @ sea level. Allowable
0.715 SG
0.715 SG
overpressure = 10% for single valve, 16% for multiple valves
PRO-RATION FACTOR APPLIED TO SIMULATION:
HEAT & MATERIAL BALANCES
PLANT:
UNIT:
EQ NO:
EQ SVC:
P&ID:
RFCC-IRVING OIL
RFCC-IRVING OIL
C-42517
DEBUTANIZER
Ax-D-425-123
PSV: 00951/00952/00954
SET PRESSURE:
185 PSIG
OVER PRESSURE:
16 %
ATM PRESSURE:
14.70 PSIA
RELIEF PRESSURE:
229.30 PSIA
STREAMS
NOTE
IN:
FEED
REBOILER
BASIS
FACTOR
1.00
1.00
TOTAL
OUT:
728,303
CASE DESCRIPTION:
PHASE
°API
M
NORMAL
MW
86.6
°F
288.1
PSIA
167.7
BTU/LB
158.26
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
TOTAL
1.00
1.00
1.00
1.00
1.00
144,908
0
583,395
728,303
MESSAGE:
BASIS:
All data obtained from SWEC Case 3.
NOTES:
LB/HR
1.00
BTU/HR
115,261,233
80,880,000
196,141,233
L
L
L
50.5
111.0
154.7
45.70
105.3
382.0
169.7
186.20
6,622,296
0
108,628,149
58,990,000
22,120,000
196,360,445
Contract :
Revision :
Date :
By:
Chk:
4499920
A
26-Jun-2000
JLH
DK
Appv'd:
DDC
Page 2 of 12
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
STREAMS
NOTE
IN:
FEED
REBOILER
BASIS
FACTOR
1,3
2
1.00
0.00
TOTAL
OUT:
728,303
CASE DESCRIPTION:
PHASE
°API
TOTAL POWER FAILURE
MW
M
86.6
°F
301.1
PSIA
232.3
BTU/LB
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1
1.00
0.00
1.00
0.00
0.00
1,3
4,6
5,6
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
6.
BTU/LB
207.09
88.85
118.24
LB/HR
-257,478
144,908
0
BTU/HR
129,142,688
0
129,142,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
30,008,998
0
129,577,863
0
0
159,586,861
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°F
k=Cp/Cv
Z
cP
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
0.0
Feed from Stripper C-42515 continues to flow to the Debutanizer until empty. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler has no heat input due to the shutdown of HCO Pumparound.
Bottoms product exits at relieving conditions. Feed temperature and enthalpy are corrected for increased heat transfer of feed/bottoms exchanger.
The condenser duty is reduced to 25% (natural convection) due to loss of fans.
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
Loss of reflux pump leads to flooding of the condensers.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
26-Jun-2000
JLH
DK
Appv'd:
DDC
Page 3 of 12
BTU/LB
BTU/HR
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
STREAMS
IN:
FEED
REBOILER
BASIS
NOTE
FACTOR
1
2
0.00
1.00
TOTAL
OUT:
LB/HR
0
M
86.6
301.1
232.3
177.32
0
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,5
1
3,5
4,5
0.00
0.00
0.00
0.00
0.00
TOTAL
0
0
0
80,880,000
80,880,000
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
0
Accumulation
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
NOTES:
Partial Power Failure - 1
CASE DESCRIPTION: Feed Stops and Reflux Pump Stops
PHASE
°API
MW
°F
PSIA
0
0
0
0
0
0
BTU/LB
207.09
88.85
118.24
LB/HR
684,032
0
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
50.50
50.50
°F
178.40
178.40
k=Cp/Cv
1.27
1.27
Z
cP
684,032
50.50
178.4
1.2735
0.77
0.77
0.01
0.01
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 stops due to stripper feed stop. Overhead and bottoms products must stop to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
The condenser duty is reduced to 62.5% due to loss of 2 of 4 fans (100% for the half with the fans running and 25% for the half with natural convection only).
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
One of the Reflux/Product pumps is lost. Assume the autostart instruments fail to start spare pump and condensers flood.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
26-Jun-2000
JLH
DK
Appv'd:
DDC
Page 4 of 12
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
STREAMS
IN:
FEED
REBOILER
BASIS
NOTE
FACTOR
1,3
2
1.00
1.00
TOTAL
OUT:
728,303
M
86.6
301.1
232.3
BTU/LB
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,6
1,3
4,6
5,6
1.00
0.00
1.00
0.00
0.00
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
6.
BTU/HR
129,142,688
80,880,000
210,022,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
Partial Power Failure - 2
CASE DESCRIPTION: Feed Continues and Reflux Pump Stops
PHASE
°API
MW
°F
PSIA
30,008,998
0
129,577,863
0
0
159,586,861
BTU/LB
207.09
88.85
118.24
LB/HR
426,555
144,908
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
50.50
50.50
°F
178.40
178.40
k=Cp/Cv
1.27
1.27
Z
cP
571,463
50.50
178.4
1.2735
0.77
0.77
0.01
0.01
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 continues to flow to the Debutanizer. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
Bottoms product exits at relieving conditions. Feed temperature and enthalpy are corrected for increased heat transfer of feed/bottoms exchanger.
The condenser duty is reduced to 62.5% due to loss of 2 of 4 fans (100% for the half with the fans running and 25% for the half with natural convection only).
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
One of the Reflux/Product pumps is lost. Assume the autostart instruments fail to start spare pump and condensers flood.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
26-Jun-2000
JLH
DK
Appv'd:
DDC
Page 5 of 12
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
STREAMS
NOTE
IN:
FEED
REBOILER
BASIS
FACTOR
1,4
2
1.00
1.00
TOTAL
OUT:
728,303
CASE DESCRIPTION:
PHASE
°API
CW FAILURE
MW
M
86.6
°F
301.1
PSIA
BTU/LB
232.3
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,3
1,4
5
6
1.00
0.00
1.00
1.00
0.00
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
6.
BTU/HR
129,142,688
80,880,000
210,022,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
30,008,998
0
129,577,863
58,990,000
0
218,576,861
BTU/LB
207.09
88.85
118.24
LB/HR
-72,346
144,908
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°F
k=Cp/Cv
Z
cP
50.50
178.40
1.27
0.77
0.01
72,562
50.50
178.4
1.2735
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 continues to flow to the Debutanizer. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
Reflux/Product pump continues.
Bottoms product exits at relieving conditions. Feed temperature and enthalpy are corrected for increased heat transfer of feed/bottoms exchanger.
Condenser fans are operational.
The trim condenser is shutdown with loss of cooling water.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
26-Jun-2000
JLH
DK
Appv'd:
DDC
Page 6 of 12
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
STREAMS
NOTE
IN:
FEED
REBOILER
BASIS
FACTOR
1,3
2
1.00
1.00
TOTAL
OUT:
728,303
CASE DESCRIPTION:
PHASE
°API
REFLUX FAILURE
MW
M
86.6
°F
301.1
PSIA
BTU/LB
232.3
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1
1,3
4
4
1.00
0.00
1.00
0.00
0.00
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
BTU/HR
129,142,688
80,880,000
210,022,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
30,008,998
0
129,577,863
0
0
159,586,861
BTU/LB
207.09
88.85
118.24
LB/HR
426,555
144,908
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
50.50
50.50
°F
178.40
178.40
k=Cp/Cv
1.27
1.27
Z
cP
571,463
50.50
178.4
1.2735
0.77
0.77
0.01
0.01
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 continues to flow to the Debutanizer. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
Bottoms product exits at normal rate, but at relieving conditions. Temperature of feed is corrected for increased heat transfer of feed/bottoms exchanger.
Loss of Reflux pump leads to flooding of the condensers.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
26-Jun-2000
JLH
DK
Appv'd:
DDC
Page 7 of 12
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
STREAMS
IN:
FEED
REBOILER
BASIS
NOTE
FACTOR
1,3
2
1.00
1.00
TOTAL
OUT:
728,303
M
86.6
301.1
232.3
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,6
1,3
4,6
5,6
1.00
0.00
1.00
0.63
0.50
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
6.
BTU/HR
129,142,688
80,880,000
210,022,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
Partial Power Failure - 3
CASE DESCRIPTION: Feed Continues and Reflux Pump Continues
PHASE
°API
MW
°F
PSIA
BTU/LB
30,008,998
0
129,577,863
37,163,700
11,060,000
207,810,561
BTU/LB
207.09
88.85
118.24
LB/HR
18,709
144,908
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
50.50
50.50
°F
178.40
178.40
k=Cp/Cv
1.27
1.27
Z
cP
163,617
50.50
178.4
1.2735
0.77
0.77
0.01
0.01
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 continues to flow to the Debutanizer. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
Bottoms product exits at relieving conditions. Feed temperature and enthalpy are corrected for increased heat transfer of feed/bottoms exchanger.
The condenser duty is reduced to 62.5% due to loss of 2 of 4 fans (100% for the half with the fans running and 25% for the half with natural convection only).
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
One of the Reflux/Product pumps is lost. Assume the autostart instrument start spare pump and condensers do not flood.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
26-Jun-2000
JLH
DK
Appv'd:
DDC
Page 8 of 12
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
Partial Power Failure -4
STREAMS
IN:
FEED
REBOILER
BASIS
NOTE
FACTOR
1
2
LB/HR
0.00
0.00
0
TOTAL
OUT:
M
86.6
301.1
232.3
BTU/HR
177.32
0
0
0
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,5
0.00
0.00
0.00
0.63
0.50
1
3,5
4,5
0
0
TOTAL
0
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
0
Accumulation
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
NOTES:
CASE DESCRIPTION: Feed Stops and Reflux Pump Continues
PHASE
°API
MW
°F
PSIA
BTU/LB
BTU/LB
207.09
88.85
118.24
LB/HR
-405,351
0
0
0
0
0
36,868,750
11,060,000
47,928,750
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°F
k=Cp/Cv
Z
cP
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
0.0
Feed from Stripper C-42515 stops due to stripper feed stop. Overhead and bottoms products must stop to maintain material balance.
Reboiler has no heat input due to the shutdown of HCO Pumparound.
The condenser duty is reduced to 62.5% due to loss of 2 of 4 fans (100% for the half with the fans running and 25% for the half with natural convection only).
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
One of the Reflux/Product pumps is lost. Assume the autostart instrument start spare pump and condensers do not flood.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
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BTU/LB
BTU/HR
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
STREAMS
NOTE
IN:
BASIS
FACTOR
FEED
REBOILER
0
TOTAL
OUT:
CASE DESCRIPTION:
PHASE
°API
LB/HR
MW
M
86.6
°F
301.1
PSIA
232.3
177.32
0
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
0
0
TOTAL
0
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
0
Accumulation
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
BTU/LB
207.09
88.85
118.24
LB/HR
0
0
0
0
0
0
0
0
0
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°F
k=Cp/Cv
Z
cP
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
0.0
MESSAGE:
BASIS:
NOTES:
0
0
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
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FEED/BOTTOMS HEAT EXCHANGER RATING
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
NOTES:
1. In order to use this spreadsheet, all blue-color values must be
entered.
2. Heat exchanger rating calc. can start with an initial guess of 1.25
value for the Qa/Qn (guess). If re-iterate appears, click on the re-iterate
button or enter Qa/Qn (next guess) value for Qa/Qn (guess) until the
solution converges . For some reasons, if "ERR" messages appear, click
on the re-iterate button or restart the calc. by entering 1.25 value for the
Qa/Qn (guess).
3. The status message is used to indicate if the solution converges.
The converging criteria is based on the 5% difference between Qa and
Qr values.
ASSUMPTION:
1. Heat transfer is linear (phase change does not occur). Other
calculation method would be considered if non-linear heat transfer
occurs.
TOWER
T out
t in
t out
HEAT
EXCHANGERS
T in
TOWER FEED
CASE
NORMAL
RELIEF
NAME: FEED
TEMPERATURE
t in
t out
°F
°F
258.0
258.0
Assumed QA, BTU/HR
Calculated QR, BTU/HR
QR/QN
(QA/QR -1)
QA/QN (guess)
QA/QN (next guess)
RESULTS:
NOTES:
288.1
301.1
FLOW
RATE
LB/HR
728,303
728,303
ENTHALPY
in
out
BTU/LB
BTU/LB
129.00
129.00
35,196,671
35,195,651 CONVERGED
1.43
0.00%
1.43
1.433
Feed Temperature =
Feed Enthalpy =
301.1 °F
177.33 BTU/LB
158.26
177.33
TOWER BOTTOMS
NAME: BOTTOMS PRODUCT
TEMPERATURE
FLOW
T in
T out
RATE
°F
°F
LB/HR
382.0
435.7
317.0
342.5
MESSAGES (NOTE 3)
583,395
583,395
LMTD
(NOTE 2)
°F
75.1
107.6
DUTY
Q
BTU/HR
24,560,000
35,195,651
NOTES
Contract :
4499920
Revision :
A
Date :
26-Jun-2000
By:
JLH
Chk:
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Page 11 of 12
REBOILER THERMAL RATING
PLANT: RFCC-IRVING OIL
EQ NO: C-42517
UNIT: RFCC-IRVING OIL
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
NOTES:
TOWER
1. In order to use this spreadsheet, all blue-color values must be entered.
2. Zero reboiler duty case is checked in order to determine whether pinch calculations are
required for the relief conditions.
3. Full normal reboiler duty case is checked in order to determine whether pinch calculations
are required for the relief conditions.
4. Reboiler pinch calc. can start with an initial guess of 0.5 value for the Qa/Qn (guess). If
re-iterate appears, click on the re-iterate button or enter Qa/Qn (next guess) value for
Qa/Qn (guess) until the solution converges. For some reasons, if "ERR" messages appear,
tout
REBOILER
click on the re-iterate button or restart the calc. by entering 0.5 value for the Qa/Qn (guess).
Tout
5. The status message is used to indicate if the solution converges. The converging criteria
is based on the 5% difference between Qa and Qr values.
ASSUMPTION:
T in
1. Reboiler pinch calculation is based on the following assumptions: (1) clean reboiler surface
tin
surface (clean U), (2) bottoms enthalpy rise is proportional to bottoms temperature rise,
and (3) correction factor (f) of 1.
BOTTOMS PRODUCT
STREAM TEMPERATURES
TOWER BOTTOMS
CASE
t in
°F
NORMAL
RELIEF
REBOILER
t out
°F
356.0
406.7
T in
°F
382.0
435.7
LMTD
T out
°F
580.0
580.0
478.0
466.0
°F
COEFF
AREA
DUTY
Uc
A
FT²
Q
BTU/HR
BTU/HR-°F-FT²
156.9
95.6
109
109
8655
8655
Does Reboiler Use Steam?
Yes if ( T in @ Relief cond. - t out @ Relief cond. ) < = 0;
Value =
Yes if U * A * ( T out @ Normal cond. - t out @ Relief cond. ) / Qn > = 1; Value =
Assumed QA, BTU/HR
90,358,738
Calculated QR, BTU/HR
90,565,627
===>
(QA/QR -1)
QA/QN (guess)
QA/QN (next guess)
Is reboiler pinch feasible?
If feasible, BASIS FACTOR =
144
0.5
no
no
MESSAGES (NOTE 5)
CONVERGED
1.12
-0.2% NEED TO PERFORM REBOILER PINCH CALC. UNTIL SOLUTION CONVERGES, Qr (KCAL/HR) = 80880000
1.12
1.1178
CONCLUSION
NOTES:
80,880,000
90,565,627
4
no
Zero Reboiler Duty Case?
Full Normal Reboiler Duty?
QR/QN
NOTES
NO
2
3
Contract :
4499920
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Date :
26-Jun-2000
By:
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Page 12 of 12
ACCUMULATOR SURGE TIME
PLANT: RFCC-IRVING OIL
EQ NO: C-42517
UNIT: RFCC-IRVING OIL
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb1-2.XLS]Normal
INPUT DATA
TOTAL POWER
FAILURE
RELIEF CASE DESCRIPTION
Partial Power
Failure - 1
Partial Power
Failure - 2
Feed Stops and
Reflux Pump
Stops
Feed Continues
and Reflux Pump
Stops
REFLUX
FAILURE
CW FAILURE
Partial Power
Failure - 3
Partial Power
Failure -4
Feed Continues
and Reflux Pump
Continues
Feed Stops and
Reflux Pump
Continues
EQUIPMENT NUMBER:
D-42520
D-42520
D-42520
D-42520
D-42520
D-42520
D-42520
EQUIPMENT NAME:
Reflux Drum
Reflux Drum
Reflux Drum
Reflux Drum
Reflux Drum
Reflux Drum
Reflux Drum
P&ID NUMBER:
Ax-D-425-128
VESSEL ORIENTATION (1=HORIZ, 2=VERT):
ACCUMULATOR DIAMETER:
Ax-D-425-128
1
Ax-D-425-128
1
Ax-D-425-128
1
Ax-D-425-128
1
Ax-D-425-128
1
Ax-D-425-128
1
1
FT
9.50
9.50
9.50
9.50
9.50
9.50
9.50
ACCUMULATOR LENGTH T/T:
FT
24.00
24.00
24.00
24.00
24.00
24.00
24.00
ACCUMULATOR LIQUID LEVEL (Note 1):
FT
5.25
5.25
5.25
5.25
5.25
5.25
BASIS NOTE
(1)
(1)
(1)
(1)
(1)
(1)
5.25
(1)
BASIS FACTOR FOR CONDENSER (AIR COOLER)
0.25
0.63
0.63
1.00
1.00
0.63
0.63
BASIS FACTOR FOR CONDENSER (WATER COOLER)
0.50
0.50
0.50
0.00
1.00
0.50
0.50
REFLUX FLOW STOPS (Y OR N)
REFLUX FLOW RATE @ NORMAL COND
Y
FT³/HR
11,582.1
OH PRODUCT FLOW STOPS (Y OR N)
Y
Y
11,582.1
Y
Y
11,582.1
Y
N
11,582.1
N
Y
11582.10
N
N
11,582.1
N
N
11,582.1
N
OH PRODUCT FLOW RATE @ NORMAL COND
FT³/HR
4,079.6
4,079.6
4,079.6
4,079.6
4,079.6
4,079.6
4,079.6
ACCUMULATION DENSITY @ RELIEF COND
LB/FT³
35.5
35.5
35.5
35.5
35.5
35.5
35.5
CONDENSER DUTY
TRIM CONDENSER DUTY
TOTAL CONDENSER DUTY
LATENT HEAT OF ACCUMULATION
VOLUMETRIC FLOW RATE, CONDENSIBLES
TOTAL VOLUMETRIC FLOW RATE
ACCUMULATOR PARTIAL VOLUME
ACCUMULATOR TOTAL VOLUME
ACCUMULATOR SURGE VOLUME
ACCUMULATOR SURGE TIME (Note 2)
BASIS:
BTU/HR
BTU/HR
BTU/HR
BTU/LB
FT³/HR
FT³/HR
FT³
FT³
FT³
MIN
ACCUMULATOR SURGE CALCULATIONS
37,163,700
37,163,700
58,990,000
11,060,000
11,060,000
0
48,223,700
48,223,700
58,990,000
118
118
118
11,482
11,482
14,046
11,482
11,482
-1,616
1,065
1,065
1,065
1,926
1,926
1,926
861
861
861
4.5
4.5
(Note 3)
58,990,000
22,120,000
81,110,000
118
19,312
15,233
1,065
1,926
861
3.4
37,163,700
11,060,000
48,223,700
118
11,482
-4,180
1,065
1,926
861
(Note 3)
37,163,700
11,060,000
48,223,700
118
11,482
-4,180
1,065
1,926
861
(Note 3)
NOTES:
1. For reflux failure, the accumulator liquid level is at the independent high alarm level. For power failure and other
cases, the accumulator liquid level is at the normal liquid level.
2. If the surge time is 10 minutes or less, there is no credit for cooling by the overhead condenser for the relief case
being considered.
3. Accumulator will not flood. Reflux and/or overhead product flow rate leaving accumulator equals or exceeds
rate of condensate accumulation.
4. Volume calculations assume elliptical 2:1 heads.
14,747,500
11,060,000
25,807,500
118
6,145
6,145
1,065
1,926
861
8.4
APPENDIX
B-2
APPENDIX B
PAGE
1
REACTOR LOOP RELIEF LOAD
DATE
06-00
FLUOR DANIEL.
FLARE SYSTEM
PROCESS MANUAL
APPENDIX B-2
REACTOR LOOP RELIEF LOAD
This example illustrates calculation methods presented in Section 4.4.
Problem Definition
VRDS reactor effluent is normally cooled by exchange with mixed feed in E-1310 A&B before entering the hot
high pressure separator (D-1320). Vapor from the separator exchanges against mixed feed in E-1320 and
then against recycle gas in E-1330. Final cooling is completed in air cooled exchangers E-1340 and E-1350
before the effluent enters the cold high pressure separator (D-1330). Vapor from this separator passes
through an amine contactor before flowing to the suction of the recycle compressor. Determine the relief load
generated during total power failure.
Required Data
Normal operating flows, cooling curves; exchanger duties, and exchanger UA values are obtained or
calculated from simulation data.
Total Power Failure Scenario
Conditions:
M/U Gas Compressor Stops
Feed Pumps Stop
Air Fan Motors Stop
Wash Water Pumps Stop
Recycle Compressor Continues to Run
Assumptions:
Partial loss of cooling in E-1310 A/B and E-1320/21
Reactor effluent composition and temperature unchanged.
Heat exchanger UA’s sufficient for 0°C approach of hot and cold outlets
25% duty credit for natural convection in E-1340 and E-1350
Constant volume operation of C-1310
Calculations
APXB2-R1.DOC
APPENDIX
B-2
APPENDIX B
PAGE
2
REACTOR LOOP RELIEF LOAD
DATE
06-00
FLUOR DANIEL.
FLARE SYSTEM
PROCESS MANUAL
Normal exchanger UA values are calculated from the exchanger duties and temperatures. On power failure,
unit feed is lost and the air cooler fans stop. The recycle gas compressor continues to run, at least for a
while. Effluent exchange is recalculated using normal exchanger UA values and assuming that the reactor
effluent continues at its normal rate and temperature. Interchanger duties are recalculated assuming there is
no liquid feed on the cold side. Air cooler duties are reduced to 25% percent of normal. Based on the
calculated effluent temperature at the cold high pressure separator, the new volumetric vapor rate is
calculated and adjusted to relief pressure. Assuming the recycle compressor operates at constant inlet
volume flow, the relief load is calculated as the difference between the new inlet volumetric flow (at relief
pressure) and the normal operating volumetric flow. Relief vapor properties are based on the relief condition
effluent properties. Calculation summaries follow:
@ Relief
CHPS Vapor
@ Normal
MW
4.67
5.40
3
2071
2541
23.25
22.08
46.5
189
Act m /h
kg/m
3
°C
3
3
(2541 - 2071) m /h = 470 m /h
Vapor Relief Load:
3
3
1.1 x 470 m /h x 22.08 kg/ m =
APXB2-R1.DOC
11, 475 kg/h
APPENDIX
B-2
APPENDIX B
PAGE
3
REACTOR LOOP RELIEF LOAD
DATE
06-00
FLUOR DANIEL.
FLARE SYSTEM
PROCESS MANUAL
REACTION SECTION PFD
REACTOR SECTION FLOWS AND PROPERTIES
APXB2-R1.DOC
APPENDIX
B-2
APPENDIX B
PAGE
4
REACTOR LOOP RELIEF LOAD
DATE
06-00
FLUOR DANIEL.
FLARE SYSTEM
PROCESS MANUAL
REACTOR SECTION FLOWS AND PROPERTIES
APXB2-R1.DOC
APPENDIX
B-3
APPENDIX B
PAGE
1
RUPTURE TUBE RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
APPENDIX B-3
RUPTURED TUBE RELIEF LOAD
This example illustrates calculation methods presented in Section 4.7.
Problem Definition
Calculate the tower relief load resulting from a ruptured tube in the steam reboiler of a Toluene Column
system.
Required Data
1. The data required to calculate the steam flow through the tube rupture is the tube inside diameter (1 inch),
which is obtained from the exchanger data sheet, and the I.P. steam pressure and temperature which are
obtained from utility system design basis data.
2. The data required to calculate the column relief load resulting from a tube rupture is similar to the
requirement described in Appendix B-1, pages 5 & 6.
Calculations
1. Steam Flow through a Ruptured Tube:
Steam flow through the tube rupture is calculated based on equations provided in Section 4.7, using the
attached spreadsheet (Fig. B3.1). Total steam flow is calculated by the program to be about 3,150 kg/h.
If not associated with a column, this flow can be used as the relief load for that case, other wise follow
step 2.
2. Column Rupture Tube Relief Load:
The column relief load is estimated conservatively based on normal reboiler and condenser duties with an
added heat imbalance resulting from a ruptured tube. This load can be estimated as follow:
•
Follow the Sample Calculation steps in Appendix B-1 for Tower Relief Load calculation.
•
Create a spreadsheet for Tube Rupture Case (Fig. B3.2).
•
FEED BF = 1.00, REBOILER BF = 1.00, add a line for STEAM FROM TUBE RUPTURE to the “IN”
o
section, with 3,150 kg/h, and an enthalpy of vapor of 667 kcal/kg at 203 C (I.P. steam temperature).
Make sure the spreadsheet calculates the total enthalpy for this line.
APXB3-R1.DOC
Figure B3.1
Contract :
04411700
Revision :
A
Date :
15-Jun-2000
By:
Chk:
Your name here
Appv'd:
HEAT EXCHANGER TUBE RUPTURE
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb3-1.xls]Sheet1
GENERAL
PLANT:
NAC
PSV:
UNIT:
SET PRES:
EQUIP. NO.:
OVER PRES:
EQUIP. SVC.:
TOLUENE COLUMN REBOILER
P&ID NO.:
FLARE
C-SV1011 A/B
RELIEF VALVE
3.5 KG/CM²G
10 %
ATM PRES:
1.03 KG/CM²A
RELIEF PRES:
4.9 KG/CM²A
INPUT DATA (NOTE 1)
FLUID PHASE: 1=LIQUID, 2=VAPOR, OR 3=TWO PHASE (NOTE 2)
2
TUBE I.D., d
0.782 IN
DESIGN PRESSURE, HIGH PRESSURE SIDE
20 KG/CM²G
NORMAL OPERATING PRESSURE, HIGH PRESSURE SIDE, P1
15.8 KG/CM²G
NORMAL OPERATING TEMPERATURE, HIGH PRESSURE SIDE
203 °C
WEIGHT % FLASH, R (NOTE 3)
100 %
BUBBLE POINT PRESSURE, Pbp (NOTE 4)
15.8 KG/CM²A
LIQUID SG AT FLOWING CONDITIONS
1.000
VAPOR MW
18.02
VAPOR COMPRESSIBILITY, Z
HIGH PRESSURE
SIDE
HEAT EXCHANGER
1
VAPOR SPECIFIC HEAT RATIO, k
1.30
ORIFICE COEFFICIENT, C
LOW PRESSURE
SIDE
0.60
SINGLE PHASE RESULTS
API RP-521 PRESSURE RATIO TWO-THIRDS RULE:
L.P. Side @ 4.9 KG/CM²A < 2/3*H.P. Side @ 14.0 KG/CM²G
TUBE AREA
3.10 CM²
EXPANSION COEFFICIENT, Y
1. The following basis is used to determine the required relieving flow rate: the tube failure is a sharp break in one tube;
the tube failure occurs at the back side of the tube sheet; high pressure fluid flows both through the tube stub remaining
in the tube sheet and through the other, longer section of the tube. For simplification, the calculation equations are
CRITICAL HYDRAULIC PRESSURE, Pcf
9.2 KG/CM²A
VAPOR FLOW REGIME
CRITICAL
VAPOR DENSITY AT OPERATING CONDITIONS, HIGH PRESSURE SIDE, pv
2. Liquid flow is calculated based on incompressible flow assuming an orifice coefficient of 0.6; Vapor flow at critical
7.51 KG/M³
LIQUID DENSITY AT OPERATING CONDITIONS, HIGH PRESSURE SIDE, pv
TWO PHASE RESULTS
CRITICAL BUBBLE POINT HYDRAULIC PRESSURE
DELTA PRESSURE FOR TWO PHASE FLOW
LIQUID RELIEF FLOW RATE COEFFICIENT, CL
conditions is calculated using the API RP-520 equation with an orifice coefficient of 0.6 and a beta ratio of less than 0.2;
Vapor flow at sub-critical conditions uses a more generalized form of the equation; Two phase flow uses the
SINGLE PHASE LIQUID RELIEF FLOW RATE AT PSV INLET CONDITIONS, Wl
SINGLE PHASE VAPOR RELIEF FLOW RATE AT PSV INLET CONDITIONS, Wv
conservatively based on flow through two orifices.
methodology recommended in API RP-520 for sizing two-phase relief valves.
3,148 KG/HR
3. Flash the high pressure side fluid at the greater of the critical bubble point hydraulic pressure or the low pressure side
pressure to determine the weight percent of the fluid that is vapor. Alternatively, to simplify calculations,
a conservative solution can be reached by assuming no flashing occurs across the ruptured tube. Actual relief valve
sizing will be based on flashing at relief conditions.
4. If the high pressure side fluid is all liquid at normal conditions, but flashes as it passes through the orifice, calculate the
VAPOR RELIEF FLOW RATE COEFFICIENT, CV
bubble point pressure at the normal temperature of the high pressure side fluid. If the high pressure side fluid is two
AREA, LIQUID PHASE
phase at normal conditions, then enter the high pressure side normal operating pressure as the bubble point pressure.
AREA, VAPOR PHASE
TWO PHASE LIQUID RELIEF FLOW RATE
TWO PHASE VAPOR RELIEF FLOW RATE AT PSV INLET CONDITIONS
Contract :
Revision :
Date :
By:
Chk:
4499920
A
15-Jun-2000
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DK
Appv'd:
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Page 2 of 12
Figure B3.2
TOWER RELIEF LOAD CALCULATIONS
PLANT: NAC
UNIT: SULFOLANE
#VALUE!
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb3-2.xls]FIG B3.2
STREAMS
NOTE
IN:
BASIS
FACTOR
FEED
REBOILER
STEAM FROM TUBE RUPTURE
1.00
1.00
TOTAL
OUT:
KG/HR
CASE DESCRIPTION:
PHASE
°API
55,218
L
3,148
V
REBOILER TUBE RUPTURE
MW
30.7
°C
184.0
KG/CM²A
KCAL/KG
4.9
203.0
84.20
667.00
58,366
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
WATER CONDENSED
1.00
0.00
1.00
1.00
43,922
KCAL/HR
4,649,356
7,891,800
2,099,716
14,640,872
11,295
V
L
L
3,148
L
30.8
175.4
4.9
155.83
30.7
208.6
5.4
99.39
175.4
177.00
6,844,365
0
1,122,610
8,247,900
557,196
30.8
TOTAL
58,365
Accumulation
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
16,772,071
°KCAL/KG
155.83
79.16
76.67
KG/HR
-27,797
43,922
MW
50.50
°C
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°C
k=Cp/Cv
Z
cP
50.50
178.40
1.27
0.77
0.01
16,125
50.50
178.4
1.2735
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
MESSAGE:
BASIS:
NOTES:
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
APPENDIX
B-4
APPENDIX B
PAGE
1
FIRE RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
APPENDIX B-4
FIRE RELIEF LOAD
This example illustrates calculation methods presented in Section 4.8.
Problem Definition
Determine the fire case relief load for relief valve PSV00951/00952/00954. These valves protect the
Debutanizer Column system described in Appendix B-1.
Required Data
All the data required for the calculations are provided in the fire case spreadsheet. Required data includes
equipment dimensions from the P&ID or equivalent drawings, equipment elevation from the P&ID, equipment
data or plot plan, liquid levels from the P&ID and/or based on guidelines presented in Section 4.8, and latent
heat/vapor properties from the process simulation files. Equipment insulation is identified on the P&ID.
Calculations
Calculations on the attached spreadsheet (Fig. B4.1) are based on the criteria defined in Section 4.8 using
equations provided in API RP 521, for process systems with drainage. Although some of the equipment in
this system is insulated, no credit is taken for this calculation in order to minimize insulation designated as
fireproof. The Debutanizer column, with a skirt height of 19 ft, will easily have liquid exposure up to the
maximum exposed height of 25 ft based on the normal liquid level. Exposed liquid level is then the level of
maximum fire exposure (25 ft) less the 19 ft skirt height, or 6 ft. Both exchanger services (E-42537 and E42538) diameters are less than 6 ft, and are assumed to be liquid full and at grade with the full length and
diameter exposed to the fire. The condenser and accumulator are not included since they are located above
the maximum exposed height of 25 ft. In all cases, no insulation credits are taken (F=1.0). Latent heat is
equal to 120.8 btu/lb. Other vapor stream properties are shown. No load calculation is required for the air
cooled condenser per the guidelines.
Conclusions
APXB4-R1.DOC
APPENDIX
B-4
APPENDIX B
PAGE
2
FIRE RELIEF LOAD
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
The total fire case relief load calculated from the spreadsheet is about 121,278 lb/h. Since the relieving vapor
has similar properties to the tower case relief loads discussed in Appendix B-1, it is clear by inspection that
the fire case relief load is less than the maximum power failure relief load, and is not the controlling case.
APXB4-R1.DOC
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Figure B4.1
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FIRE CASE RELIEF LOAD CALCULATIONS - WETTED SURFACES
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb4-1.xls]Fire Liquid
GENERAL
PROCESS: RFCCU
PSV: PSV00951/00952/00954
UNIT: RFCCU
SET PRES:
EQUIP. NO.: C42517 (Debutanizer)
NLL
Normal
Liquid Level
If vessel diameter is
less than 1.83 meters
(6 feet), use full
surface area of vessel
as wetted surface area
- for horizontal vessels
only.
NLL
OVER PRES:
EQUIP. SVC.:
ATM PRES:
P&ID NO.: Ax-D-425-123
RELIEF PRES:
185.00 PSIG
16 %
14.70 PSI
229.30 PSIA
NOTES
1. The Debutanizer Column skirt height = 19 ft. The fractionation tower normal liquid level is
Fire
Surface
Elevation
25 ft or 7.62 m from grade (ref: Fire Case Relief Load Calculation Guidelines). Therefore, the
7.62 meters (25 feet) per API 521
Vessel
Bottom
Tangent
fire exposed liquid level for the Depropanizer is 6 ft.
2. E-42536 is the Debutanizer Condenser, therefore it will drain into the reflux drum and it
is not included.
Wetted liquid level is the wetted level up to 7.62 meters (25 feet) elevation above the fire
surface elevation. For horizontal vessels where the bottom of the vessel is above 7.62
meters, if the vessel is less than half full, calculate the wetted area at the liquid level. If
the vessel is more than half full, calculate the wetted area at half filled.
INPUT DATA
EQUIPMENT IN FIRE ZONE TAG NO.
C42517
E-42537
E-42538
EQUIPMENT SERVICE
Debutanizer
Debutanizer
Reboiler
Debutanizer
Debutanizer
Feed / Bottoms Trim
Condenser
HX
P&ID NO. (EQUIPMENT LOCATION)
Ax-D-425-123 Ax-D-425-124 Ax-D-425-123 Ax-D-425-127
VESSEL Type(1=H;2=V;3=S)
Number of Identical Vessels
2
1
1
1
1
1
1
2
Diameter, O.D.
FT
14.00
5.31
3.67
3.90
Height or Length
FT
109.33
22.00
26.00
29.80
Normal Liquid Level
FT
10.25
5.31
3.67
3.90
Vessel Elevation to Bottom
FT
19.00
14.00
1.00
1.00
Fire Exposed Liquid Level
FT
6.00
5.31
3.67
3.90
1
2
2
2
1.0
1.0
1.0
1.0
BTU/LB
120.8
120.8
120.8
120.8
°F
173.6
173.6
173.6
173.6
50.50
50.50
50.50
50.50
0.7724
0.7724
0.7724
0.7724
1.2735
1.2735
1.2735
1.2735
2,575
2,575
2,575
2,575
Number of Wetted Heads
F Factor
BOOT
E-42566
Diameter
FT
T/T
FT
PIPING Inlet Piping, O.D.
IN
Length
FT
Outlet Piping, O.D.
IN
Length
FT
F Factor
LIQUID
API Gravity
Viscosity
cP
UOPK
Heat of Vaporization
VAPOR Temperature
MW
Compressibility
Z
k=Cp/Cv Ratio
Lower Heating Value
BTU/SCF
CALCULATED RESULTS
VESSEL Wetted Surface
FT²
499.1
433.7
331.4
801.6
---
---
---
---
---
BOOT
Wetted Surface
FT²
0.0
0.0
0.0
0.0
---
---
---
---
---
PIPING
Inlet Wetted Surface
FT²
0.0
0.0
0.0
0.0
---
---
---
---
---
Outlet Wetted Surface
FT²
0.0
0.0
0.0
0.0
---
---
---
---
---
Wetted Surface
FT²
499.1
433.7
331.4
801.6
---
---
---
---
---
Heat Absorbed
BTU/HR
3,425,522
3,053,164
2,448,559
5,723,143
---
---
---
---
---
Relief Rate
LB/HR
28,357
25,275
20,270
47,377
---
---
---
---
---
LB/HR
121,278
TOTAL
TOTAL RELIEF RATE
Temperature
°F
173.6
Viscosity
cP
0.0000
Z
0.7724
MW
Compressibility
51
k=Cp/Cv Ratio
Lower Heating Value
1.2735
BTU/SCF
2,575
<===
APPENDIX
B-5
APPENDIX B
PAGE
1
RATING OF THE HORIZONTAL
FLARE K.O. DRUM
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
APPENDIX B-5
RATING OF THE HORIZONTAL FLARE K.O. DRUM
This example illustrates the calculation method, presented in Section 8.1.2.1, used to verify the size of the
existing horizontal flare K.O. drum.
Problem Definition
1. Based on a given drum diameter, length and properties of the relief stream, determine the maximum relief
load that the existing horizontal drum can safely handle to knockout liquid droplets 400 microns in diameter.
2. Based on a given drum diameter, length and properties of the relief stream, determine the maximum liquid
droplet diameter for a relief load of 260,100 kg/hr.
Required Data
All the data required for the calculations are summarized in Table 1. The required data is divided into two
categories; flare drum dimensional data and flare gas data.
Flare gas data is primarily derived from the physical/chemical properties of the gas. Based on an evaluation of
the K.O. drum pumpout system, the K.O. drum liquid level is assumed to be 0.7 meters. The built-up
2
backpressure at the drum is calculated to be about 1. 33 Kg/cm A. Liquid density is typically assumed to be
3
700 kg/m when the specific liquid information is not available.
Calculations
Calculations are based on the sizing criteria described in API RP 521, Section 5.4 , 3rd edition, November 1990,
for a horizontal flare K.O. drum. These calculations are shown on the attached spreadsheets (Fig. B5.1 & B5.2)
for the input data provided.
Conclusions
1. The calculations (Fig. B5.1) indicate that a horizontal K.O. drum of 4.5 meter diameter and 15 meter length is
adequate to knockout liquid droplets 400 microns in diameter for a total relief load of about 197,000 kg/hr of
the relief gas of the given properties outlined above.
2. For the same K.O. drum with a total relief load of 260,100 kg/hr of the same relief gas, this spreadsheet (Fig.
B5.2) calculates the maximum removable droplet size of about 530 microns, which is in the acceptable
range of 400-600 microns for an existing drum.
APXB5-R1.DOC
APPENDIX
B-5
APPENDIX B
PAGE
2
RATING OF THE HORIZONTAL
FLARE K.O. DRUM
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
TABLE B5-1
INPUT DATA
Flare Gas Data
2
•
Atmospheric pressure:
1.03 kg/cm A
•
Drum Operating Pressure:
1.33 kg/cm A
•
Relief vapor temperature:
142.8 °C
•
Droplet size to be knocked out:
400 microns
•
Relief vapor viscosity:
0.01 cp
•
Relief vapor molecular weight:
11.8
•
Liquid Density:
700 kg/m
2
Flare K.O. Drum Dimensional Data
•
Drum diameter
4.5 m
•
Drum Length
15 m
•
Liquid Level in the drum after the relief:
0.7 m
APXB5-R1.DOC
3
Figure B5.1
Contract :
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Date :
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HORIZONTAL K.O. RATING SPREADSHEET
NRP FLARE K.O. DRUM (Y-V2907)
TOTAL POWER FAILURE CASE
DRUM DATA
Length
15 m
49.2 ft
Diameter
4.5 m
14.8 ft
Radius
2.25 m
7.38 ft
Liquid Level
0.70 m
2.3 ft
Vapor Space
3.8 m
12.5 ft
FLUID DATA
Temperature
Pressure
M.W.
Liquid Density
Gas Viscosity
Gas Density
Liq. Drop. Vel.
Liq. Drop. Time
Vap. Vel.
142.8
1.33
11.8
700
0.01
0.45
2.2
1.8
8.5
øC
kg/cmýA
kg/m
cp
kg/m
m/s
s
m/s
289.0
18.98
11.8
43.7
0.01
0.03
7.1
1.8
28.1
øF
psia
lb/ft
cp
lb/ft
ft/s
s
ft/s
PROBLEM TYPE
Capacity Known
Capacity
Particle Size
Next Guess
Converge
converge
No leave capacity blank
kg/hr
400 microns
lb/hr
0.00131 ft
399
PARAMETERS CALCULATIONS
C(Re)
2.6E+03
Drag Coeff. (C)
1.73
Tot. Flow Area
15.9 m
Area for Liq. Flow
1.6 m
Area for Vap. Flow
14.3 m
Liq. Volume
24 m
% capacity diff.
100.0
RESULTS SUMMARY
Drum Capacity
196,679 kg/hr
Particle Size
400 microns
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb5-1.xls]Sheet1
2.6E+03
1.73
171.2
17.0
154.2
835
ft
ft
ft
ft
434,379 lb/hr
0.001312336 ft
Figure B5.2
Contract :
04411700
Revision :
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Date :
15-Jun-2000
By:
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HORIZONTAL K.O. RATING SPREADSHEET
NRP FLARE K.O. DRUM (Y-V2907)
TOTAL POWER FAILURE CASE
DRUM DATA
Length
15 m
49.2 ft
Diameter
4.5 m
14.8 ft
Radius
2.25 m
7.38 ft
Liquid Level
0.70 m
2.3 ft
Vapor Space
3.8 m
12.5 ft
FLUID DATA
Temperature
Pressure
M.W.
Liquid Density
Gas Viscosity
Gas Density
Liq. Drop. Vel.
Liq. Drop. Time
Vap. Vel.
142.8
1.33
11.8
700
0.01
0.45
2.9
1.3
11.3
øC
kg/cmýA
kg/m
cp
kg/m
m/s
s
m/s
289.0
18.98
11.8
43.7
0.01
0.03
9.4
1.3
37.1
øF
psia
lb/ft
cp
lb/ft
ft/s
s
ft/s
PROBLEM TYPE
Capacity Known
Capacity
Particle Size
Yes enter capacity (kg/hr) below
260,100 kg/hr
529 microns
Next Guess
Converge
converge
573,416 lb/hr
0.00174 ft
528
PARAMETERS CALCULATIONS
C(Re)
6.0E+03
Drag Coeff. (C)
1.31
Tot. Flow Area
15.9 m
Area for Liq. Flow
1.6 m
Area for Vap. Flow
14.3 m
Liq. Volume
24 m
% capacity diff.
0.0
RESULTS SUMMARY
Drum Capacity
260,176 kg/hr
Particle Size
529 microns
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb5-2.xls]Sheet1
6.0E+03
1.31
171.2
17.0
154.2
835
ft
ft
ft
ft
574,484 lb/hr
0.001735564 ft
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
B-6
APPENDIX B
PAGE
1
FLARE RADIATION
DATE
06-00
FLARE SYSTEM
APPENDIX B-6
FLARE RADIATION
This example illustrates calculation methods presented in Section 9.4.
Problem Definition
Based on a maximum relief load of 1,120,000 kg/hr, determine if the radiation level at the plot limits established
for the future location of a process unit is acceptable.
Required Data
All the data required for the calculations are summarized in Table 1. The required data is divided into three
categories: flare gas data, flare stack dimensional data, and wind/air data.
Flare gas data is primarily derived form the physical/chemical properties of the gas. Lean flammability limits of
the relief vapor are estimated adequately by interpolation of pure component values with similar molecular weight
and characteristics (aromatic, olefinic, parafFinic). Atmospheric pressure must also be input along with flare
radiation factors. For maximum relief loads, Section 9.4 of the manual indicates that the fraction of heat radiated
(F) is 1.0 for flare radiation calculations.
Flare stack dimensional data are obtained directly from flare system data sheets and drawings. A plot plan is
used to determine the vertical and horizontal distances from the flare base to grade at the future plant boundary.
Wind velocity is set at 8.9 m/sec per the criteria defined in Section 9.4.
Calculations
Calculations are based on the criteria defined in Section 9.4, using the Brzustowski and Sommer method. These
calculations are shown on a spreadsheet format (Figure B6.1) for the input data provided. In addition to
calculating the radiation level at the future process unit boundary, input tables are also provided on the
spreadsheet which allow radiation levels to be calculated at various distances from the stack and/or distances
from the stack calculated for varying radiation levels. Each of these values is calculated based on the elevation
level input for the primary target location (future plant boundary).
APXB6-R1.DOC
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
B-6
APPENDIX B
PAGE
2
FLARE RADIATION
DATE
06-00
FLARE SYSTEM
Conclusions
2
For this case the calculations indicate that the maximum acceptable radiation level of 4054 kcal/h-m is
2
acceptable at the future plant boundary, 65 meters from the stack. A peak radiation level of 4190 kcal/h-m
10
occurs 32 meters from the stack. Total flare heat release is 1.11 x 10 kcal/h and the tip Mach No. is 0.43,
which is low enough that the flame will be easily stabilized.
APXB6-R1.DOC
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
B-6
APPENDIX B
PAGE
3
FLARE RADIATION
DATE
06-00
FLARE SYSTEM
TABLE B6-1
INPUT DATA
Flare Gas Data
2
•
Atmospheric pressure:
1.03 kg/cm A
•
Relief vapor temperature:
115 °C
•
Relief flow rate:
1,120,000 kg/h
•
Relief vapor Cp/Cv:
1.08
•
Relief vapor net heating value:
9,899 kcal/kg
•
Relief vapor lean flammability limit:
2.09%
•
Relief vapor molecular weight:
48.9
•
Fraction of heat radiated:
•
Fraction of heat transmitted:
.15
1.0
Flare Stack Dimensional Data
•
Flare tip diameter
1.5 m
•
Flare stack height
128.0 m
•
Vertical distance from stack base to object:
0.0 m
•
Horizontal distance from stack base to object:
65.0 m
Wind/Air Data
•
Wind velocity:
8.9 m/s
•
Air molecular weight
28.96
•
Air temperature
21.1 °C
APXB6-R1.DOC
FIGURE B6.1
Contract :
04411700
Revision :
A
X_l
Date :
15-Jun-2000
By:
FLARE RADIATION CALCULATIONS (B&S METHOD)
Your name here
Chk:
Appv'd:
WIND
L
S_l
NAC FLARE STACK (MAXIMUM CAPACITY: 1,120,000 KG/HR)
NOTES:
θ
Z_l
1. Brzustowski's and Sommer method for flare radiation calculation is taken from API Recommended,
Practice, 3rd ed., Nov. 1990.
2. Ideal Gas Law is used.
Xm - H
3. Lean flammability limit is taken from pure components values adjusted for the mol. wt. of flared gas.
Yc
XXcc
RADIATION INTENSITY (K) VERSUS HORIZONTAL DISTANCE FROM STACK BASE TO OBJECT (R)
K==========> R Table
d
R==========> K Table
K
R
K
R
R
K
R
K
kcal/hr/m
m
Btu/hr/ft
ft
m
kcal/hr/m
ft
Btu/hr/ft
1,400
283
500
949
0
4,055
0
1,495
2,700
164
1,000
537
50
4,150
100
1,545
4,100
59
1,500
207
65
4,054
200
1,506
8,200
#NUM!
3,000
#NUM!
100
3,661
213
1,495
13,600
#NUM!
5,000
#NUM!
150
2,914
300
1,391
Rmin = Xc
Kmax
m
kcal/hr/m
32
PEAK RADIATION INTENSITY
4,190
Rmin = Xc
Kmax
ft
Btu/hr/ft
106
Flare Gas Data
P
Absolute Pressure
Tg
Temperature
W
Flow Rate
k
NHV
D
H
H' = H + Yc
N = R - Xc
1,545
R
Metric Units
1.03 kg/cm^2(A)
Specific Heat Ratio = Cp/Cv
115.00 °C
239.00 °F
2,469,152 lb/hr
9,899 kcal/kg
CL
Lean flammability Limit (vol % or mol %): enter fraction value
Mg
Molecular Weight
0.02
48.90 g/gmol
Note
14.70 psia
1,120,000 kg/hr
1.08
Net Heating Value
English Units
1.08
17,818 Btu/lb
0.02
3
48.90 lb/lbmol
F
Fraction of Heat Recieved as Radiant Heat
0.15
0.15
t
Fraction of Radiant Heat Transmitted
1.00
1.00
Flare Stack Dimensional Data
d
Flare Tip Diameter
1.5 m
59.25 in
H
Flare Stack Height
128.0 m
419.95 ft
h
Vertical Distance from Stack Base to Object (negative if object below stack base)
R
Horizontal Distance from Stack Base to Object
0.0 m
0.00 ft
65.0 m
213.25 ft
Ua
Wind Velocity
Ma
Air Molecular Weight
Ta
Air Temperature
21.1 °C
ρ
Air Density (kg/m3) = Ma*P*/(0.084784*(Ta + 273.16)).
1.20 kg/m
0.07 lb/ft
2
1.54 kg/m
0.10 lb/ft
2
Wind/Air Data
8.9 m/sec
28.96
29.20 ft/sec
28.96
69.98 °F
3
Air Density (lb/ft ) = Ma*P/(10.732*(Ta + 459.67))
Flare Gas Calculations
ρ
Gas Density (kgm3) = Mg*P*/(0.084784*(Tg + 273.16))
Gas Density (lb/ft3) = Mg*P*/(10.732*(Tg + 459.67))
V
Volumetric Gas Flow Rate = W/ gas density/3600
203 mü/sec
7,157 ftü/sec
Ug
Gas Exit Velocity = 4V/(pi*d^2) [m/sec] or 4*144V/(pi*d^2) [ft/sec]
114 m/sec
374 ft/sec
Us
Gas Sonic Velocity = 91.2(k*(Tg+273.16)/Mg)^0.5 [m/sec] or 223(k*(Tg+459.67)/Mg)^0.5 [ft/sec]
267 m/sec
876 ft/sec
Ug/Us
0.43
Mach #
dP
Pressure Drop at Flare Tip (m H2O) = (8.27*10^-7*W*((Tg+273.16)/Mg)^0.5/d^2)^2
1.3 m H2O
0.43
52.3 in H2O
Pressure Drop at Flare Tip (in H2O)= (2.72*10^-3*W*((Tg+459.67)/Mg)^0.5/d^2)^2
Q
Heat Released = W * NHV
1.11E+10 kcal/hr
4.40E+10 Btu/hr
Qr
Heat Transmitted = Q * F (emissivity)
1.66E+09 kcal/hr
6.60E+09 Btu/hr
Dimensionaless Parameters Calculations
C_L
CL*(Ug/Ua)*(Mg/Ma)
45.2%
45.2%
S_L
If C_L <=0.5, S_L = 2.04/(C_L^1.03). If C_L > 0.5, S_L = 2.51/(C_L^0.625)
4.63
4.63
X_L
If C_L <=0.5, X_L = S_L - 1.65. If C_L > 0.5 and S_L > 2.35, X_L = S_L - 1.65.
2.98
2.98
If C_L > 0.5 and S_L <= 2.35, X_L = 10^((log(S_L) - log (1.04) - log (2.05))/2.28)
Z_L
2.05*X_L^0.28
R_
(Ug/Ua)*(Gas density/Air density)^0.5
2.78
2.78
14.48
14.48
Z_l
Z_L * d * R_L [m] or Z_L* d * R_L/12 [ft]
60.6 m
X_l
X_L * d * R_L [m] or X_L * d * R_L/12 [ft]
64.8 m
212.7 ft
Xc
X_l/2
32.4 m
106.4 ft
Yc
Z_I * 0.82
49.7 m
163.1 ft
Dimensional Parameters Calculations
198.9 ft
Flame Parameters Calculations
q
Angle of Flare Flame from Vertical = ATAN (Xc/Yc)
33.1
L
Flame Length = SQRT(Z_l^2 + X_l^2)
88.8 m
291.2 ft
100.8 m
330.7 ft
S_l
Actual Curved Lengh of Flame = S_L * d * R_L [m] or S_L * d * R_L/12 [ft]
33.1
Single Point Radiation Parameters Calculations
N
Horizontal Distance from Flame Radiation Source Point to Object = |R - Xc|
H'
Vertical Distance from Object to the Flame Radiation Source Point = |H + Yc - h|
D
Diagonal Distance from Flame Radiation Source Point to Object = (N^2 + H'^2)^0.5
K
Radiation Intensity = (Tau * Qr)/(4pi *D^2)
H:\SP\CENLIBRY\P000\225\Vol 48\[Figb6-1.xls]FIG B6.1
32.6 m
106.9 ft
177.7 m
583.1 ft
180.7 m
4,054 kcal/hr/m
592.8 ft
1,495 Btu/hr/ft
APPENDIX
B-7
APPENDIX B
PAGE
1
DYNAMIC SIMULATION
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
APPENDIX B-7
DYNAMIC SIMULATION
This example illustrates calculation methods presented in Section 5.3.6.
Problem Definition
The No. 4 Crude Column was identified as a good candidate for dynamic simulation due to its large
contribution to the flare load. The objective of the study was to determine the relief associated with a total
power failure and to provide a better understanding of the complex behavior in the Crude Column during the
relief event. Table 1 provides the individual effects and the key assumptions which defined the basis for the
simulation.
Required Data
All of the data required to perform the dynamic simulation of the No. 4 Crude Column are summarized in
Table 2. Detailed mechanical information is necessary in order to develop a realistic model of the system. A
steady state simulation from PRO II or Hysim is used as a reference and validation tool.
Calculations
The envelope for the dynamic simulation began at the Crude Flash Drum and included all downstream heat
exchange. Since the ID and FD fans for the Crude Heater shutdown during a total power failure, and
safeguarding functions initiate fuel shut-off in the event of a power failure, a duty decay curve derived from
mechanical analysis of performance and design data for the fired heater was used to set the heat absorbed by
the process. The Crude Column was represented with 21 trays. Mass and energy hold-ups were factored to
properly account for the inventory in the actual column. The reflux, KERO, LGO, and HGO pump arounds
were modeled as single duties. The KERO and LGO Strippers were not modeled. However, the dynamic
model did account for the flows to and from these vessels. A simple model of the HGO Stripper was
developed. The model of the overhead equipment included the air-cooled exchanger, water cooler, and the
receiver. Refer to Diagram 1 for a sketch of the system.
APXB7-R1.DOC
APPENDIX
B-7
APPENDIX B
PAGE
2
DYNAMIC SIMULATION
DATE
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FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
The simulation model was validated at steady state conditions with a Simsci Process simulation. The crude
feed was recharacterized with a reduced number of pseudo-components for computational efficiency.
After the dynamic model was validated at steady state, transient analysis began. In the transient analysis, the
system was disturbed (flows set to zero, cooling duties lost, etc.) and the response was tested. As defined in
Table 1, two cases were investigated. In Case 1, the reflux to the column failed. Case 2 examined the effect
of continued reflux on the peak relief load. Time dependent profiles for pressures, temperatures, flowrates,
duties, etc. for each case were generated and analyzed (refer to Figures 1 through 15).
Conclusions
Based on the attached graphical results, the total power failure relief load with failure of the reflux is
approximately 470,000 kg/hr. With reflux continuing, the peak relief was reduced to 255,000 kg/hr.
APXB7-R1.DOC
APPENDIX
B-7
APPENDIX B
PAGE
3
DYNAMIC SIMULATION
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
TABLE B7-1: Crude Column Dynamic Simulation Problem Definition
Relief Case:
Total power failure.
Individual Effects:
Pumps -
P-4001 (Crude Charge Pump):
One motor and one turbine driven pump are normally operating. The
motor driven pump fails. The crude feed continues at 50% of the
normal rate. Because the simulation envelope begins at the Crude
Flash drum (V-4007), this assumption affects only cases where the
reflux pump remains operational and heat exchange in E-4001
continues. Cases with reflux failure are not affected by this
assumption.
P-4002 (Desalted Crude Pump):
Normally operating pump is turbine driven. Normal flow from the
Crude Flash Drum continues.
P-4003 (Red. Crude R/D Pump):
Normally operating pumps are turbine driven. Normal flow from the
bottom of the Crude Column continues. Heat exchange in E-4006,
4008, 4013, 4014, and 4015 continues.
P-4004 (Circulating HGO Pump):
Normally operating pump is motor driven, as is the spare. Pump
fails, HGO pump-around is lost. Heat exchange in E-4012, 4018,
and 4029 stops.
P-4005 (Circulating LGO Pump):
Normally operating pumps are motor driven, as is the spare. Pumps
fail, LGO pump-around is lost. Heat exchange in E-4011 and E-4028
stops.
APXB7-R1.DOC
APPENDIX
B-7
APPENDIX B
PAGE
4
DYNAMIC SIMULATION
DATE
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FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
P-4006 (Circ. KERO Pump):
Normally operating pump is motor driven, as is the spare. Pump
fails, KERO pump-around is lost. Heat exchange in E-4005 and E4007 stops.
P-4007 (HGO R/D Pump):
Normally operating pump is motor driven, as is the spare. Pump
fails, heat exchange in E-4004 and 4016 stops.
P-4008 (LGO R/D Pump):
Normally operating pump is motor driven, as is the spare. Pump
fails, heat exchange in E-4003, E-4010 and E-4019 stops.
P-4009 (KERO R/D Pump):
Normally operating pump is motor driven, as is the spare. Pump
fails, heat exchange in E-4002, E-4009, and E-4017 stops.
P-4010 (Crude Reflux Pump):
Normally operating pump is motor driven. Spare pump is turbine
driven with autostart. The dynamic simulation will investigate two
cases:
Case 1: Motor driven pump is in operation. Pump fails, as does
autostart.
Case 2: Turbine driven pump is in operation. Reflux continues.
P-4011 (LGO Strip. O/H Pump):
Normally operating pump is motor driven, as is the spare. Pump fails,
LGO return to the Crude Column stops.
P-4016 (Splitter Charge Pump):
Normally operating pumps are motor driven, as is the spare. Pumps
fails, crude column overhead to the Splitter stops.
P-4029 (O/H Condensate Pump):
Normally operating pump is motor driven, as is the spare. Pump
fails.
P-4030 (Acc. Recov. Oil Pump):
Normally operating pump is motor driven, as is the spare. Pump
fails.
Heat Exchangers -
APXB7-R1.DOC
APPENDIX
B-7
APPENDIX B
PAGE
5
DYNAMIC SIMULATION
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
E-4021 (LGO Flash Vap. Cooler):
Loss of cooling water. Duty set to zero.
E-4022 (Overhead Fan Cooler):
Loss of all fans. Duty set to 40% for natural convection.
E-4023 (Overhead Trim Cooler):
Loss of cooling water. Duty set to zero.
E-4041 (Col. Reflux Fan Cooler):
Loss of all fans. In Case 2, duty will be set to 40% for natural
convection.
E-4001 (Col. Reflux / Crude):
In Case 2, the duty of E-4001 will be set to 50% of normal to account
for the reduction of fresh crude flow on the tube side due to the
failure of P-4001A.
Fired Heater -
H-4001 (Crude Heater): In the event of a power failure, the following two safeguarding functions can initiate
the shutdown of the Crude Heater:
1. The 6.6 kV low voltage relays
2. The low air pressure switch PSLL-4041 A/B. Air pressure would fall due
to the loss of the forced draft fans.
Upon loss of flame, the heater duty will drop sharply to a small percentage of
the normal duty. The duty then decays from this point as the refractory cools.
This behavior was simulated using a duty decay curve derived from analysis
of the heater mechanical design and performance data.
Miscellaneous The flow of stripping steam to V-4001 and V-4002 continues. The flow will be dependent on the pressure in
the columns.
Summary of Cases:
APXB7-R1.DOC
APPENDIX
B-7
APPENDIX B
PAGE
6
DYNAMIC SIMULATION
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Two cases will be studied. The differences between these cases are summarized below.
Case 1:
Crude Reflux Pump (P-4010) fails.
Case 2:
Crude Reflux Pump (P-4010) remains in operation.
Key Modeling Assumptions:
The starting point for the dynamic simulation will be the Crude Flash Drum (V-4007). The flow of liquid and
flash vapor from V-4007 will be assumed to remain constant for the duration of the events. Since this drum
has a large hold-up time (in excess of 10 minutes), crude from the drum will also be assumed to be at a
constant temperature. This is a conservative assumption because most of the heat exchange upstream of V4007 will be lost during a total power failure, and, therefore, the crude temperature will gradually decrease.
APXB7-R1.DOC
APPENDIX
B-7
APPENDIX B
PAGE
7
DYNAMIC SIMULATION
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
TABLE B7-2: Required Information
General Information
1.
PFD's and P&ID's
2.
Detailed Steady State Simulation
Mechanical Information
1.
2.
2.
3.
Exchanger Data:
•
Surface Area
•
Tube Information (I.D., thickness, number, material)
•
Design Duty
•
Surface Area (Bare and Extended)
•
Tube Information (I.D., thickness, number, material)
•
Fin information (O.D., pitch, thickness, material)
•
Design Duty
•
Air Flow
•
Air Inlet and Outlet Temperature
•
Fan Information (power and number)
Air-cooled Exchanger Data:
Fired Heater Data and Mechanical Drawings
•
Surface Area
•
Refractory Design and Material
•
Tube Information (I.D., thickness, number)
•
Dimensions (I.D., T/T)
•
General Assembly Drawings
Column Data
APXB7-R1.DOC
APPENDIX
B-7
APPENDIX B
PAGE
8
DYNAMIC SIMULATION
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
APXB7-R1.DOC
•
Tray Detail Drawings
•
Normal Liquid Levels
APPENDIX
B-7
APPENDIX B
PAGE
9
DYNAMIC SIMULATION
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
The following figures are attached for the Case 1:
1. PSV Flow
2. PSV Molecular Weight
3. Crude Column Feeds
4. Pumparound Flows
5. Side Draw Flows
6. Overhead Flow
7. Bottoms Flow
8. Column Pressure Profile
9. H-4001 Outlet Temperature
10. Column Temperature Profile
11. H-4001 Duty
12. Heat Integration Duties
13. Pump Around Duties
14. Condensing Duties
15. Overhead Receiver Level
APXB7-R1.DOC
APPENDIX
B-8
APPENDIX B
PAGE
1
FLARE SYSTEM HYDRAULICS
CALCULATIONS
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
APPENDIX B-8
FLARE SYSTEM HYDRAULICS CALCULATIONS
This example illustrates calculation methods presented in Section 7.3.
Problem Definition
Plant power failure results in the HOU flare relief loads summarized in the attached load summary (Table 8.1).
Determine the impact of header backpressure on relief valve capacity.
Required Data
In addition to the relief rate required from individual relief valves, definition of the complete hydraulic network is
required, including the flare stack, flare header, and laterals. Relief valve set pressures and sizes are also
required.
Calculations
The complete hydraulic network, including laterals for the relieving valves, is defined and input into the INPLANT
program. Required individual relief valve flows from the load summary are also input and the resultant
backpressure at the outlet of each relieving valve calculated. This procedure is outlined in the attached summary
entitled “Flare System Hydraulics Evaluation Activities”. Related source data required for the calculations is also
attached.
Conclusions
Backpressures calculated at the outlet of each relief valve are shown in the attached HOU Flare Header
Hydraulics Sketch (Figure 8.5). The attached one-page HOU Flare Hydraulics summary highlights any
backpressures calculated in excess of 10% of relief valve set pressure for conventional relief valves and 30% of
set pressure for balanced bellows valves. In addition to the total power failure case, results from other major
power failure cases are also provided in the summary.
For the total power failure case, these backpressure values are exceeded for only two relief valves. The value of
11 percent for the Debutanizer relief valve is slightly in excess of the targeted maximum of 10 percent for
conventional valves, but would be considered marginally acceptable. However, as shown in the summary, a
much higher backpressure (17 percent) is calculated for another failure mode considered, which requires the
relief valve to be replaced with a balanced bellows valve more tolerant to backpressure.
APXB8-R1.DOC
APPENDIX
B-8
APPENDIX B
PAGE
2
FLARE SYSTEM HYDRAULICS
CALCULATIONS
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
The 41 percent backpressure calculated for the balanced bellows Atmospheric Fractionator relief valve results in
only about a one percent reduction in valve maximum capacity and is acceptable in this case. As indicated in the
summary, backpressure for this relief valve is even higher (43 percent) for another case. Flow capacity of the
relief valve is determined to be adequate for that case after accounting for the reduction in flow capacity due to
the high backpressure. Therefore, no valve change is required.
APXB8-R1.DOC
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
C
APPENDIX C
PAGE
1
PSV SIZING SOFTWARE
DATE
06-00
FLARE SYSTEM
APPENDIX C
PSV SIZING SOFTWARE
The majority of the major relief valve manufacturers have developed software to assist is sizing the type of
relief valves that they provide. This software is generally available free on request by potential purchasers.
These programs can be used to assist in valve sizing. The purchased valve selection may be different from
this early trial selection and should be established only after discussion with a qualified representative of the
relief valve supplier. Crosby, Consolidated and Anderson, Greenwood & Company have provided permission
to use their programs and copies of those programs to Yukong Limited. A description of the capabilities of
these programs follows:
CROSBY
CROSBY-SIZE v2.1, 21 Oct 94 (CSIZE.BAT)
CROSBY-SIZE, version 2.1, is an MS-Dos program that can run on a stand alone personal
computer, be shared on a network. It will run under MS-Windows . Crosby-Size presents
recommendations for ASME Section I and VIII valve selection based on Crosby’s experience in sizing
and selection of safety relief and pressure relief devices. Use of this program allows an accurate
determination of such parameters as orifice size, maximum flow and noise. The program provides
calculations, user selected units, selection of valve size and style, valve data storage and printed
reports and specification sheets.
Crosby-Size has the option to calculate the amount of vapor generated in a vessel under fire
conditions in accordance with the guidelines published in API RP 520.
A 92 page User’s Manual is available with instructions to install and operate this program
Important Note: The current version of Crosby-Size is “date stamped”. It will only function until 01
Aug 97. After that date you will need to contact Crosby Valve to obtain a new version of the program.
A typical example of the output of the Crosby-Size program has been included as Figure C.1.
CROSBY BOILER SET SIZE v1.0, 03 Apr 95 (BSIZE.BAT)
An ASME Section I safety valve is a device designed to protect power boilers during an overpressure
event.
The proper sizing, selection, manufacturing, assembly, testing and maintenance are all critical to
obtain optimum protection.
This program is designed to provide recommendations based on Crosby’s experience in sizing,
selecting, testing, installing and operating safety relief and pressure relief valves. The program
provides supplemental technical information such as reaction forces (total resultant force), noise
levels and general dimensional information on the valve selection.
A 50 page User’s manual is available with instructions to install and operate this program.
APXC1-R1.DOC
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
C
APPENDIX C
PAGE
2
PSV SIZING SOFTWARE
DATE
06-00
FLARE SYSTEM
Important Note: The current version of Crosby Boiler Set Size is “date stamped”. It will only function
until 03 April 98. After that date you will need to contact Crosby Valve to obtain a new version of the
program.
Note: After installing Boiler Set Size run the program PCXDATL.EXE to unzip the data files.
Full distribution rights for both programs (CSIZE.BAT and BSIZE.BAT) are granted with the following
restrictions:
1.
2.
3.
The Crosby-Size System must be distributed in it’s original form.
No charge can be extracted, except for nominal disk copying fees.
No company, person, or persons may claim ownership rights to it other than Crosby
Valve and Gage Company.
Technical problems are supported at the following location:
Crosby Valve & Gage Company
43 Kendrick Street
Wrentham, MA 02093
Telephone: (508) 384-3121
Fax: (508) 384-8675
CONSOLIDATED 
SRVS4 Version 4.1 (DRESSER.EXE)
Consolidated  Sizing & Selection Engineering Program - SRVS4
SRVS Version 4.0 is a computer program that is to be used to size and select Consolidated  Safety
Relief Valves. The program was developed with a Graphical Interface Utility which provides menus,
lists and forms for easy data input. The program also will detect if a pointer device (mouse) is
installed and will allow the use of the device. In addition to the sizing and selection of safety relief
valves, the program can create a database to store the valve selection data for future use and allow
user specification of fluid data and material compatibility.
This program is provided with a 37 page operating manual. When ordering, it is advisable to request
a copy of the SRV-1 catalog. The SFV-1 catalog provides specific relief valve information, sizing
equations w/appropriate corrections, flange pressure/temperature ratings, valve dimensions, and
instructions for proper installation.
This program allows a data base for individual PSVs and another data base for fluid properties. A
fluid physical properties data base can be developed for those streams that are frequently used,
saving the effort of having to enter the data again and again.
A typical example of the Consolidated program output has been included in this appendix as Figure
C.2.
Technical problems are supported at the following location:
Dresser Valve & Controls Division
Post Office Box 1430
Alexandria, Louisiana 71309-1430
APXC1-R1.DOC
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
C
APPENDIX C
PAGE
3
PSV SIZING SOFTWARE
DATE
06-00
FLARE SYSTEM
Phone: (318) 640-2250 Telex: 58-6423 Cable: DIVID
APXC1-R1.DOC
Fax: (318) 640-6222
APPENDIX
C
APPENDIX C
PAGE
4
PSV SIZING SOFTWARE
DATE
06-00
FLARE SYSTEM
FLUOR DANIEL
.
PROCESS MANUAL
ANDERSON, GREENWOOD & CO
SIZING.EXE, Version 3.0, Feb 94
The SIZING.EXE program can be used for both ASME Section I and Section VIII relief valve sizing.
An Intel 386 or higher based computer, with 4M of hard disk space available for installation, is
required. As valve files are added to the database, more hard disk space is required.
Use of the Anderson, Greenwood Co.’s software is available free of charge and may be freely copied.
Their only restriction is that the software should be registered. This will assure updates and notice of
any changes or corrections.
Version 3.1 is anticipated to be distributed in mid-January of 1996. It is supported by Microsoft
Windows. Ms. Strauss can be contacted for details at the address/phone given below.
A typical example of the Anderson, Greenwood & Co program output has been included in this
appendix as Figure C.3.
SUBSONIC, Version 3.0
The SUBSONIC.EXE program is available for relief valves with pressure below 15 PSIG or about
1.0 kg/cm2 G.
Technical problems are supported at the following location:
Anderson, Greenwood & Co.
Attn: Michele Strauss
P.O. Box 944
Stafford, Tx 77497
Phone: (713) 274-4511
Fax: (713) 274-4644
PASSWORD
The SIZING.EXE program was developed with a “freeze” date of 31 Dec 94. At this time a password
was required to run the program and is asked for when the program is opened. The correct password
is:
SALES
until 31 Dec 95
REVENUES
beginning 1 Jan 96
APXC1-R1.DOC
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
C
APPENDIX C
PAGE
5
PSV SIZING SOFTWARE
DATE
06-00
FLARE SYSTEM
FIGURE C.1
TYPICAL CROSBY PRV REPORT SHEET
APXC1-R1.DOC
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
C
APPENDIX C
PAGE
6
PSV SIZING SOFTWARE
DATE
06-00
FLARE SYSTEM
FIGURE C.2
TYPICAL CONSOLIDATED SRV SIZING REPORT
APXC1-R1.DOC
FLUOR DANIEL
.
PROCESS MANUAL
APPENDIX
C
APPENDIX C
PAGE
7
PSV SIZING SOFTWARE
DATE
06-00
FLARE SYSTEM
FIGURE C.3
TYPICAL ANDERSON, GREENWOOD
& COMPANY PRV REPORT
APXC1-R1.DOC
APPEJDIX
D-1
APPENDIX D
PAGE
1
FLARE SYSTEM CALCULATION
SPREADSHEETS
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
APPENDIX D-1
FLARE SYSTEM CALCULATION SPREADSHEETS
Included in this Appendix are hard copies of Microsoft Excel Spreadsheets which can be used for
flare system calculations. The associated spreadsheet software can be found under Flare System
Calculation Spreadsheets in the Fluor Daniel Intranet/Functions/Engineering/Disciplines/
Process/Aliso Viejo/Process Knowledge/ Process Design Manuals. Following are brief descriptions of
these spreadsheets. All input data required is color-coded in blue.
1.
Tower Load Calculation
An example of the use of the Tower Load Calculation spreadsheet is provided in Appendix B-1.
The Tower Load Calculation is based on defining a heat and material balance envelop around a
tower. General information, the unit design basis, tower design basis, and normal properties and flow
rates are entered on the first sheet, shown as Figure D1.1. Up to eight different relief cases can be
defined on the next eight spreadsheets, Figures D1.2 through 9. The first seven sheets have sample
relief cases, e.g., Total Power Failure, 4 Partial Power Failures, Cooling Water Failure, Reflux failure,
which can be changed by changing the case description title.
The Heat Exchanger Rating for Relief Load Calculation, in Figure D1.10, estimates the tower feed
temperature and enthalpy at relief conditions by re-rating the tower feed-bottoms exchanger,
assuming a constant value for UA. Enter the input values, which are color-coded in blue, and follow
the messages and notes displayed in the sheet, pressing the Re-Iterate button as required until the
program converges. Go to the relief case where Heat Exchanger re-rating applies and enter the new
feed temperature and enthalpy.
The Reboiler Thermal Rating for Relief Load Calculation, in Figure D1.11, estimates a re-rated
reboiler duty at relief conditions assuming a constant value for UA. Enter the input values, colorcoded in blue, and follow the messages and notes displayed in the sheet, pressing the Re-Iterate
button as required until the program converges. Go to the relief case where reboiler thermal rating
applies and adjust the Basis Factor until the reboiler duty in the relief case equals the Reboiler
Thermal Rating relief duty.
Accumulator Surge Time, shown in Figure D1.12, uses a simple mass balance and condensation rate
to estimate the surge time for up to eight relief cases. Again all input data required is color-coded in
blue.
APXD1-R1.DOC
2.
APPEJDIX
D-1
APPENDIX D
PAGE
2
FLARE SYSTEM CALCULATION
SPREADSHEETS
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
Liquid and Vapor Fire Case Relief Load Calculation
Both Liquid and Vapor Fire Case calculations are based on methodology provided in API RP-521, and
discussed in Section 4.8. Up to nine separate pieces of equipment and fluid properties can be
entered in each sheet. Provisions to account for identical vessels have also been provided, as well as
vessel boots and associated piping. The Vapor Fire Case calculation determines a relief load based
on ideal gas behavior. Other properties input to this spreadsheet are for use in the Pressure Safety
(Relief) Valve Specification Sheet. Liquid and Vapor Fire Case spreadsheets are shown in Figures
D1.13 and 14, respectively. A sample calculation using the Liquid Fire Case Relief Load Calculation
is provided in Appendix B-4.
3.
Heat Exchanger Tube Rupture
The Heat Exchanger Tube Rupture spreadsheet uses the method discussed in Section 4.7 to
determine the flow rate passing through a tube rupture at relief conditions on the downstream side.
The relief valve load must then be calculated separately, based on considerations discussed in
Section 4.7. The spreadsheet, shown in Figure D1.15, calculates flow rate from all liquid, all vapor, or
two phase streams under non-critical and critical conditions. A sample calculation using this
spreadsheet is provided in Appendix B-3.
4.
Flare Radiation Calculation
The flare stack height or the radiation intensity is calculated using the B&S, Simple, and Kent
methods as shown in Figures D1.18 through 20. The standard calculation method is B&S and is
discussed in Section 9.4.4. Enter the input data in English units or metric units, Figures D1.16 and
17, respectively, and then press the macro button at the top center of the input sheet. The LHV
calculations are based on average molecular weight determined from the major relief loads, per
Figure D1.21. A sample calculation using this sheet is provided in Appendix B-6.
5.
Flare Horizontal Knock Out Drum Rating
The Flare Horizontal Knock Out Drum Rating spreadsheet, Figure D1.22, calculates either particle
dropout size for a fixed flow rate or flow rate for a required droplet size. The calculation uses the API
sizing method described in Section 8.1. Appendix B-5 provides a sample calculation using this sheet.
APXD1-R1.DOC
6.
APPEJDIX
D-1
APPENDIX D
PAGE
3
FLARE SYSTEM CALCULATION
SPREADSHEETS
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
PSV Sizing Data Sheet and User’s Guide
Figures D1.23 through 25 show the User’s Guide and PSV Sizing Data Sheets, respectively.
The User’s Guide provides an over view and step-by-step walkthrough of the spreadsheet.
The PSV Sizing Data Sheet (Figure D1.24) is based on sizing methods provided in API RP-520 and
discussed in Section 3.11 and valve orifice sizes from API Standard 526 (shown in Figure 3.13). This
data sheet provides preliminary valve sizing which is useful for process engineers during estimating
and early design phases. Sizing equations are for preliminary check sizing only. All sizes are to
be confirmed by a vendor. Input areas are color-coded blue. Fill in the General section, equipment
design and operating conditions and backpressure. Enter the data for each relief condition for all
liquid, all vapor, or two phase relief loads. The weight percent flash is based on the inlet percent
liquid. The program will not calculate the flashed relief load. You must flash the relief stream and
enter the resultant liquid and vapor portions and properties, manually. If the valve being specified has
liquid trim, press the Menu button and select liquid trim.
When issuing a process data sheet to control system, Figure D1.25 should be used. This is the same
process PSV data sheet as in Figure D1.24 with the preliminary valve sizing section hidden.
7.
Pressure Safety (Relief) Valve Specification Sheet
Use the Pressure Safety (Relief) Valve Specification Sheet, shown in Figure D1.26, to write sizing
specifications and manufacturer information if the existing data sheet is not available. This
specification sheet is usually provided by control system at the end of each project.
APXD1-R1.DOC
Contract :
4499920
Revision :
Date :
By:
Chk:
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 1 of 12
Figure D1.1
TOWER RELIEF LOAD CALCULATIONS
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
GENERAL
TOWER DESIGN BASIS
UNIT FEED RATE
TOWER FEED RATE
70,000 BPD @ 60°F
70,000 BPD @ 60°F
464 KL/H @ 60°F
464 KL/H @ 60°F
7,952 T/D
7,952 T/D
66.5 °API
66.5 °API
Relief Pressure = Set Pressure + Allowable Overpressure + Atmospheric Pressure (14.696 @ sea level. Allowable
0.715 SG
0.715 SG
overpressure = 10% for single valve, 16% for multiple valves
PRO-RATION FACTOR APPLIED TO SIMULATION:
HEAT & MATERIAL BALANCES
PLANT:
UNIT:
EQ NO:
EQ SVC:
P&ID:
RFCC-IRVING OIL
RFCC-IRVING OIL
C-42517
DEBUTANIZER
Ax-D-425-123
PSV: 00951/00952/00954
SET PRESSURE:
185 PSIG
OVER PRESSURE:
16 %
ATM PRESSURE:
14.70 PSIA
RELIEF PRESSURE:
229.30 PSIA
STREAMS
NOTE
IN:
FEED
REBOILER
BASIS
FACTOR
1.00
1.00
TOTAL
OUT:
728,303
CASE DESCRIPTION:
PHASE
°API
M
NORMAL
MW
86.6
°F
288.1
PSIA
167.7
BTU/LB
158.26
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
TOTAL
1.00
1.00
1.00
1.00
1.00
144,908
0
583,395
728,303
MESSAGE:
BASIS:
All data obtained from SWEC Case 3.
NOTES:
LB/HR
1.00
BTU/HR
115,261,233
80,880,000
196,141,233
L
L
L
50.5
111.0
154.7
45.70
105.3
382.0
169.7
186.20
6,622,296
0
108,628,149
58,990,000
22,120,000
196,360,445
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 2 of 12
Figure D1.2
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
STREAMS
NOTE
IN:
FEED
REBOILER
BASIS
FACTOR
1,3
2
1.00
0.00
TOTAL
OUT:
728,303
CASE DESCRIPTION:
PHASE
°API
TOTAL POWER FAILURE
MW
M
86.6
°F
301.1
PSIA
232.3
BTU/LB
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1
1.00
0.00
1.00
0.00
0.00
1,3
4,6
5,6
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
6.
BTU/LB
207.09
88.85
118.24
LB/HR
-257,478
144,908
0
BTU/HR
129,142,688
0
129,142,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
30,008,998
0
129,577,863
0
0
159,586,861
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°F
k=Cp/Cv
Z
cP
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
0.0
Feed from Stripper C-42515 continues to flow to the Debutanizer until empty. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler has no heat input due to the shutdown of HCO Pumparound.
Bottoms product exits at relieving conditions. Feed temperature and enthalpy are corrected for increased heat transfer of feed/bottoms exchanger.
The condenser duty is reduced to 25% (natural convection) due to loss of fans.
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
Loss of reflux pump leads to flooding of the condensers.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 3 of 12
BTU/LB
BTU/HR
Figure D1.3
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
STREAMS
IN:
FEED
REBOILER
BASIS
NOTE
FACTOR
1
2
0.00
1.00
TOTAL
OUT:
LB/HR
0
M
86.6
301.1
232.3
177.32
0
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,5
1
3,5
4,5
0.00
0.00
0.00
0.00
0.00
TOTAL
0
0
0
80,880,000
80,880,000
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
0
Accumulation
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
NOTES:
Partial Power Failure - 1
CASE DESCRIPTION: Feed Stops and Reflux Pump Stops
PHASE
°API
MW
°F
PSIA
0
0
0
0
0
0
BTU/LB
207.09
88.85
118.24
LB/HR
684,032
0
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
50.50
50.50
°F
178.40
178.40
k=Cp/Cv
1.27
1.27
Z
cP
684,032
50.50
178.4
1.2735
0.77
0.77
0.01
0.01
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 stops due to stripper feed stop. Overhead and bottoms products must stop to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
The condenser duty is reduced to 62.5% due to loss of 2 of 4 fans (100% for the half with the fans running and 25% for the half with natural convection only).
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
One of the Reflux/Product pumps is lost. Assume the autostart instruments fail to start spare pump and condensers flood.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 4 of 12
Figure D1.4
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
STREAMS
IN:
FEED
REBOILER
BASIS
NOTE
FACTOR
1,3
2
1.00
1.00
TOTAL
OUT:
728,303
M
86.6
301.1
232.3
BTU/LB
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,6
1,3
4,6
5,6
1.00
0.00
1.00
0.00
0.00
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
6.
BTU/HR
129,142,688
80,880,000
210,022,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
Partial Power Failure - 2
CASE DESCRIPTION: Feed Continues and Reflux Pump Stops
PHASE
°API
MW
°F
PSIA
30,008,998
0
129,577,863
0
0
159,586,861
BTU/LB
207.09
88.85
118.24
LB/HR
426,555
144,908
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
50.50
50.50
°F
178.40
178.40
k=Cp/Cv
1.27
1.27
Z
cP
571,463
50.50
178.4
1.2735
0.77
0.77
0.01
0.01
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 continues to flow to the Debutanizer. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
Bottoms product exits at relieving conditions. Feed temperature and enthalpy are corrected for increased heat transfer of feed/bottoms exchanger.
The condenser duty is reduced to 62.5% due to loss of 2 of 4 fans (100% for the half with the fans running and 25% for the half with natural convection only).
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
One of the Reflux/Product pumps is lost. Assume the autostart instruments fail to start spare pump and condensers flood.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 5 of 12
Figure D1.5
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
STREAMS
NOTE
IN:
FEED
REBOILER
BASIS
FACTOR
1,4
2
1.00
1.00
TOTAL
OUT:
728,303
CASE DESCRIPTION:
PHASE
°API
CW FAILURE
MW
M
86.6
°F
301.1
PSIA
BTU/LB
232.3
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,3
1,4
5
6
1.00
0.00
1.00
1.00
0.00
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
6.
BTU/HR
129,142,688
80,880,000
210,022,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
30,008,998
0
129,577,863
58,990,000
0
218,576,861
BTU/LB
207.09
88.85
118.24
LB/HR
-72,346
144,908
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°F
k=Cp/Cv
Z
cP
50.50
178.40
1.27
0.77
0.01
72,562
50.50
178.4
1.2735
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 continues to flow to the Debutanizer. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
Reflux/Product pump continues.
Bottoms product exits at relieving conditions. Feed temperature and enthalpy are corrected for increased heat transfer of feed/bottoms exchanger.
Condenser fans are operational.
The trim condenser is shutdown with loss of cooling water.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 6 of 12
Figure D1.6
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
STREAMS
NOTE
IN:
FEED
REBOILER
BASIS
FACTOR
1,3
2
1.00
1.00
TOTAL
OUT:
728,303
CASE DESCRIPTION:
PHASE
°API
REFLUX FAILURE
MW
M
86.6
°F
301.1
PSIA
BTU/LB
232.3
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1
1,3
4
4
1.00
0.00
1.00
0.00
0.00
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
BTU/HR
129,142,688
80,880,000
210,022,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
30,008,998
0
129,577,863
0
0
159,586,861
BTU/LB
207.09
88.85
118.24
LB/HR
426,555
144,908
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
50.50
50.50
°F
178.40
178.40
k=Cp/Cv
1.27
1.27
Z
cP
571,463
50.50
178.4
1.2735
0.77
0.77
0.01
0.01
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 continues to flow to the Debutanizer. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
Bottoms product exits at normal rate, but at relieving conditions. Temperature of feed is corrected for increased heat transfer of feed/bottoms exchanger.
Loss of Reflux pump leads to flooding of the condensers.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 7 of 12
Figure D1.7
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
STREAMS
IN:
FEED
REBOILER
BASIS
NOTE
FACTOR
1,3
2
1.00
1.00
TOTAL
OUT:
728,303
M
86.6
301.1
232.3
177.32
728,303
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,6
1,3
4,6
5,6
1.00
0.00
1.00
0.63
0.50
TOTAL
144,908
583,395
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
6.
BTU/HR
129,142,688
80,880,000
210,022,688
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
728,303
Accumulation
NOTES:
LB/HR
Partial Power Failure - 3
CASE DESCRIPTION: Feed Continues and Reflux Pump Continues
PHASE
°API
MW
°F
PSIA
BTU/LB
30,008,998
0
129,577,863
37,163,700
11,060,000
207,810,561
BTU/LB
207.09
88.85
118.24
LB/HR
18,709
144,908
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
50.50
50.50
°F
178.40
178.40
k=Cp/Cv
1.27
1.27
Z
cP
163,617
50.50
178.4
1.2735
0.77
0.77
0.01
0.01
0.7724
0.0105
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
Feed from Stripper C-42515 continues to flow to the Debutanizer. Overhead and bottoms products must continue flowing to maintain material balance.
Reboiler operates normally with HCO Pumparound operational.
Bottoms product exits at relieving conditions. Feed temperature and enthalpy are corrected for increased heat transfer of feed/bottoms exchanger.
The condenser duty is reduced to 62.5% due to loss of 2 of 4 fans (100% for the half with the fans running and 25% for the half with natural convection only).
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
One of the Reflux/Product pumps is lost. Assume the autostart instrument start spare pump and condensers do not flood.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 8 of 12
Figure D1.8
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
Partial Power Failure -4
STREAMS
IN:
FEED
REBOILER
BASIS
NOTE
FACTOR
1
2
LB/HR
0.00
0.00
0
TOTAL
OUT:
M
86.6
301.1
232.3
BTU/HR
177.32
0
0
0
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
1,5
0.00
0.00
0.00
0.63
0.50
1
3,5
4,5
0
0
TOTAL
0
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
0
Accumulation
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
MESSAGE:
BASIS:
1.
2.
3.
4.
5.
NOTES:
CASE DESCRIPTION: Feed Stops and Reflux Pump Continues
PHASE
°API
MW
°F
PSIA
BTU/LB
BTU/LB
207.09
88.85
118.24
LB/HR
-405,351
0
0
0
0
0
36,868,750
11,060,000
47,928,750
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°F
k=Cp/Cv
Z
cP
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
0.0
Feed from Stripper C-42515 stops due to stripper feed stop. Overhead and bottoms products must stop to maintain material balance.
Reboiler has no heat input due to the shutdown of HCO Pumparound.
The condenser duty is reduced to 62.5% due to loss of 2 of 4 fans (100% for the half with the fans running and 25% for the half with natural convection only).
The trim condenser duty is reduced to 50% due to loss of motor driven cooling water pump.
One of the Reflux/Product pumps is lost. Assume the autostart instrument start spare pump and condensers do not flood.
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 9 of 12
BTU/LB
BTU/HR
Figure D1.9
TOWER RELIEF LOAD CALCULATIONS
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
STREAMS
NOTE
IN:
BASIS
FACTOR
FEED
REBOILER
0
TOTAL
OUT:
CASE DESCRIPTION:
PHASE
°API
LB/HR
MW
M
86.6
°F
301.1
PSIA
232.3
177.32
0
OVHD PRODUCT
SOUR WATER
BOTTOMS PRODUCT
CONDENSER
TRIM CONDENSER
0
0
TOTAL
0
V
L
L
50.5
178.4
229.3
207.09
105.3
435.7
234.3
222.11
0
Accumulation
Top Tray Vapor
Top Tray Liquid
Latent Heat of Vaporization
RELIEF LOAD:
ACCUMULATION (Note A)
OVHD PRODUCT (Note A)
OFF GAS
STEAM
===> RELIEF LOAD
BTU/LB
207.09
88.85
118.24
LB/HR
0
0
0
0
0
0
0
0
0
MW
50.50
°F
178.4
k=Cp/Cv
1.2735
Z
0.7724
cP
0.0105
MW
°F
k=Cp/Cv
Z
cP
Accumulation =
(Heat In)-(Heat Out)
Accumulation Latent Heat of Vaporization
0.0
MESSAGE:
BASIS:
NOTES:
0
0
A. If the accumulation relief rate plus overhead product relief rate is negative, then the properties for the accumulation and overhead product are not entered.
Contract :
Revision :
Date :
By:
Chk:
4499920
A
23-Jun-2000
JLH
DK
Appv'd:
DDC
Page 10 of 12
Figure D1.10
FEED/BOTTOMS HEAT EXCHANGER RATING
PLANT: RFCC-IRVING OIL
UNIT: RFCC-IRVING OIL
EQ NO: C-42517
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
NOTES:
1. In order to use this spreadsheet, all blue-color values must be
entered.
2. Heat exchanger rating calc. can start with an initial guess of 1.25
value for the Qa/Qn (guess). If re-iterate appears, click on the re-iterate
button or enter Qa/Qn (next guess) value for Qa/Qn (guess) until the
solution converges . For some reasons, if "ERR" messages appear, click
on the re-iterate button or restart the calc. by entering 1.25 value for the
Qa/Qn (guess).
3. The status message is used to indicate if the solution converges.
The converging criteria is based on the 5% difference between Qa and
Qr values.
ASSUMPTION:
1. Heat transfer is linear (phase change does not occur). Other
calculation method would be considered if non-linear heat transfer
occurs.
TOWER
T out
t in
t out
HEAT
EXCHANGERS
T in
TOWER FEED
CASE
NORMAL
RELIEF
NAME: FEED
TEMPERATURE
t in
t out
°F
°F
258.0
258.0
Assumed QA, BTU/HR
Calculated QR, BTU/HR
QR/QN
(QA/QR -1)
QA/QN (guess)
QA/QN (next guess)
RESULTS:
NOTES:
288.1
301.1
FLOW
RATE
LB/HR
728,303
728,303
ENTHALPY
in
out
BTU/LB
BTU/LB
129.00
129.00
35,196,671
35,195,651 CONVERGED
1.43
0.00%
1.43
1.433
Feed Temperature =
Feed Enthalpy =
301.1 °F
177.33 BTU/LB
158.26
177.33
TOWER BOTTOMS
NAME: BOTTOMS PRODUCT
TEMPERATURE
FLOW
T in
T out
RATE
°F
°F
LB/HR
382.0
435.7
317.0
342.5
MESSAGES (NOTE 3)
583,395
583,395
LMTD
(NOTE 2)
°F
75.1
107.6
DUTY
Q
BTU/HR
24,560,000
35,195,651
NOTES
Contract :
4499920
Revision :
A
Date :
23-Jun-2000
By:
JLH
Chk:
DK
Appv'd:
DDC
Page 11 of 12
Figure D1.11
REBOILER THERMAL RATING
PLANT: RFCC-IRVING OIL
EQ NO: C-42517
UNIT: RFCC-IRVING OIL
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
NOTES:
TOWER
1. In order to use this spreadsheet, all blue-color values must be entered.
2. Zero reboiler duty case is checked in order to determine whether pinch calculations are
required for the relief conditions.
3. Full normal reboiler duty case is checked in order to determine whether pinch calculations
are required for the relief conditions.
4. Reboiler pinch calc. can start with an initial guess of 0.5 value for the Qa/Qn (guess). If
re-iterate appears, click on the re-iterate button or enter Qa/Qn (next guess) value for
Qa/Qn (guess) until the solution converges. For some reasons, if "ERR" messages appear,
tout
REBOILER
click on the re-iterate button or restart the calc. by entering 0.5 value for the Qa/Qn (guess).
Tout
5. The status message is used to indicate if the solution converges. The converging criteria
is based on the 5% difference between Qa and Qr values.
ASSUMPTION:
T in
1. Reboiler pinch calculation is based on the following assumptions: (1) clean reboiler surface
tin
surface (clean U), (2) bottoms enthalpy rise is proportional to bottoms temperature rise,
and (3) correction factor (f) of 1.
BOTTOMS PRODUCT
STREAM TEMPERATURES
TOWER BOTTOMS
CASE
t in
°F
NORMAL
RELIEF
REBOILER
t out
°F
356.0
406.7
T in
°F
382.0
435.7
LMTD
T out
°F
580.0
580.0
478.0
466.0
°F
COEFF
AREA
DUTY
Uc
A
FT²
Q
BTU/HR
BTU/HR-°F-FT²
156.9
95.6
109
109
8655
8655
Does Reboiler Use Steam?
Yes if ( T in @ Relief cond. - t out @ Relief cond. ) < = 0;
Value =
Yes if U * A * ( T out @ Normal cond. - t out @ Relief cond. ) / Qn > = 1; Value =
Assumed QA, BTU/HR
90,358,738
Calculated QR, BTU/HR
90,565,627
===>
(QA/QR -1)
QA/QN (guess)
QA/QN (next guess)
Is reboiler pinch feasible?
If feasible, BASIS FACTOR =
144
0.5
no
no
MESSAGES (NOTE 5)
CONVERGED
1.12
-0.2% NEED TO PERFORM REBOILER PINCH CALC. UNTIL SOLUTION CONVERGES, Qr (KCAL/HR) = 80880000
1.12
1.1178
CONCLUSION
NOTES:
80,880,000
90,565,627
4
no
Zero Reboiler Duty Case?
Full Normal Reboiler Duty?
QR/QN
NOTES
NO
2
3
Contract :
4499920
Revision :
A
Date :
23-Jun-2000
By:
JLH
Chk:
DK
Appv'd:
DDC
Page 12 of 12
Figure D1.12
ACCUMULATOR SURGE TIME
PLANT: RFCC-IRVING OIL
EQ NO: C-42517
UNIT: RFCC-IRVING OIL
PSV: 00951/00952/00954
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-1-12.xls]Dialog2
INPUT DATA
TOTAL POWER
FAILURE
RELIEF CASE DESCRIPTION
EQUIPMENT NUMBER:
D-42520
Partial Power
Failure - 1
Partial Power
Failure - 2
Feed Stops and
Reflux Pump
Stops
Feed Continues
and Reflux Pump
Stops
D-42520
D-42520
REFLUX
FAILURE
CW FAILURE
D-42520
D-42520
Partial Power
Failure - 3
Partial Power
Failure -4
Feed Continues
and Reflux Pump
Continues
Feed Stops and
Reflux Pump
Continues
D-42520
D-42520
EQUIPMENT NAME:
Reflux Drum
Reflux Drum
Reflux Drum
Reflux Drum
Reflux Drum
Reflux Drum
Reflux Drum
P&ID NUMBER:
Ax-D-425-128
Ax-D-425-128
Ax-D-425-128
Ax-D-425-128
Ax-D-425-128
Ax-D-425-128
Ax-D-425-128
VESSEL ORIENTATION (1=HORIZ, 2=VERT):
ACCUMULATOR DIAMETER:
1
1
1
1
1
1
1
FT
9.50
9.50
9.50
9.50
9.50
9.50
9.50
ACCUMULATOR LENGTH T/T:
FT
24.00
24.00
24.00
24.00
24.00
24.00
24.00
ACCUMULATOR LIQUID LEVEL (Note 1):
FT
5.25
5.25
5.25
5.25
5.25
5.25
BASIS NOTE
(1)
(1)
(1)
(1)
(1)
(1)
5.25
(1)
BASIS FACTOR FOR CONDENSER (AIR COOLER)
0.25
0.63
0.63
1.00
1.00
0.63
0.63
BASIS FACTOR FOR CONDENSER (WATER COOLER)
0.50
0.50
0.50
0.00
1.00
0.50
0.50
REFLUX FLOW STOPS (Y OR N)
REFLUX FLOW RATE @ NORMAL COND
Y
FT³/HR
OH PRODUCT FLOW STOPS (Y OR N)
11,582.1
Y
Y
11,582.1
Y
11,582.1
Y
4,079.6
N
11,582.1
N
Y
11582.10
N
OH PRODUCT FLOW RATE @ NORMAL COND
FT³/HR
4,079.6
ACCUMULATION DENSITY @ RELIEF COND
LB/FT³
35.5
CONDENSER DUTY
TRIM CONDENSER DUTY
TOTAL CONDENSER DUTY
LATENT HEAT OF ACCUMULATION
VOLUMETRIC FLOW RATE, CONDENSIBLES
TOTAL VOLUMETRIC FLOW RATE
ACCUMULATOR PARTIAL VOLUME
ACCUMULATOR TOTAL VOLUME
ACCUMULATOR SURGE VOLUME
ACCUMULATOR SURGE TIME (Note 2)
BASIS:
BTU/HR
BTU/HR
BTU/HR
BTU/LB
FT³/HR
FT³/HR
FT³
FT³
FT³
MIN
NOTES:
1. For reflux failure, the accumulator liquid level is at the independent high alarm level. For power failure and other
cases, the accumulator liquid level is at the normal liquid level.
2. If the surge time is 10 minutes or less, there is no credit for cooling by the overhead condenser for the relief case
being considered.
3. Accumulator will not flood. Reflux and/or overhead product flow rate leaving accumulator equals or exceeds
rate of condensate accumulation.
4. Volume calculations assume elliptical 2:1 heads.
14,747,500
11,060,000
25,807,500
118
6,145
6,145
1,065
1,926
861
8.4
4,079.6
Y
N
11,582.1
N
N
11,582.1
N
4,079.6
4,079.6
4,079.6
4,079.6
35.5
35.5
35.5
ACCUMULATOR SURGE CALCULATIONS
37,163,700
37,163,700
58,990,000
11,060,000
11,060,000
0
48,223,700
48,223,700
58,990,000
118
118
118
11,482
11,482
14,046
11,482
11,482
-1,616
1,065
1,065
1,065
1,926
1,926
1,926
861
861
861
4.5
4.5
(Note 3)
35.5
35.5
35.5
58,990,000
22,120,000
81,110,000
118
19,312
15,233
1,065
1,926
861
3.4
37,163,700
11,060,000
48,223,700
118
11,482
-4,180
1,065
1,926
861
(Note 3)
37,163,700
11,060,000
48,223,700
118
11,482
-4,180
1,065
1,926
861
(Note 3)
Contract :
04411700
Revision :
A
Date :
23-Jun-2000
By:
Figure D1.13
Your Name Here
Chk:
Appv'd:
FIRE CASE RELIEF LOAD CALCULATIONS - WETTED SURFACES
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
GENERAL
PROCESS: RFCCU
PSV: PSV00951/00952/00954
UNIT: RFCCU
SET PRES:
EQUIP. NO.: C42517 (Debutanizer)
NLL
Normal
LiquidLevel
Ifvesseldiameteris
lessthan1.83meters
(6feet),usefull
surfaceareaofvessel
aswettedsurfacearea
-forhorizontalvessels
only.
NLL
OVER PRES:
EQUIP. SVC.:
ATM PRES:
P&ID NO.: Ax-D-425-123
RELIEF PRES:
185.00 PSIG
16 %
14.70 PSI
229.30 PSIA
NOTES
1. The Debutanizer Column skirt height = 19 ft. The fractionation tower normal liquid level is
25 ft or 7.62 m from grade (ref: Fire Case Relief Load Calculation Guidelines). Therefore, the
Vessel
Bottom
Tangent
Fire
Surface
Elevation
7.62meters(25feet)perAPI521
fire exposed liquid level for the Depropanizer is 6 ft.
2. E-42536 is the Debutanizer Condenser, therefore it will drain into the reflux drum and it
is not included.
Wettedliquidlevelisthewettedlevelupto7.62meters(25feet)elevationabovethefire
surfaceelevation. Forhorizontalvesselswherethebottomofthevesselisabove7.62
meters,ifthevesselislessthanhalffull,calculatethewettedareaattheliquidlevel. If
thevesselismorethanhalffull,calculatethewettedareaathalffilled.
INPUT DATA
EQUIPMENT IN FIRE ZONE TAG NO.
C42517
E-42537
E-42538
EQUIPMENT SERVICE
Debutanizer
Debutanizer
Reboiler
Debutanizer
Debutanizer
Feed / Bottoms Trim
Condenser
HX
P&ID NO. (EQUIPMENT LOCATION)
Ax-D-425-123 Ax-D-425-124 Ax-D-425-123 Ax-D-425-127
VESSEL Type(1=H;2=V;3=S)
Number of Identical Vessels
2
1
1
1
1
1
1
2
Diameter, O.D.
FT
14.00
5.31
3.67
3.90
Height or Length
FT
109.33
22.00
26.00
29.80
Normal Liquid Level
FT
10.25
5.31
3.67
3.90
Vessel Elevation to Bottom
FT
19.00
14.00
1.00
1.00
Fire Exposed Liquid Level
FT
6.00
5.31
3.67
3.90
1
2
2
2
1.0
1.0
1.0
1.0
BTU/LB
120.8
120.8
120.8
120.8
°F
173.6
173.6
173.6
173.6
50.50
50.50
50.50
50.50
0.7724
0.7724
0.7724
0.7724
1.2735
1.2735
1.2735
1.2735
2,575
2,575
2,575
2,575
Number of Wetted Heads
F Factor
BOOT
E-42566
Diameter
FT
T/T
FT
PIPING Inlet Piping, O.D.
IN
Length
FT
Outlet Piping, O.D.
IN
Length
FT
F Factor
LIQUID
API Gravity
Viscosity
cP
UOPK
Heat of Vaporization
VAPOR Temperature
MW
Compressibility
Z
k=Cp/Cv Ratio
Lower Heating Value
BTU/SCF
CALCULATED RESULTS
VESSEL Wetted Surface
FT²
499.1
433.7
331.4
801.6
---
---
---
---
---
BOOT
Wetted Surface
FT²
0.0
0.0
0.0
0.0
---
---
---
---
---
PIPING
Inlet Wetted Surface
FT²
0.0
0.0
0.0
0.0
---
---
---
---
---
Outlet Wetted Surface
FT²
0.0
0.0
0.0
0.0
---
---
---
---
---
Wetted Surface
FT²
499.1
433.7
331.4
801.6
---
---
---
---
---
Heat Absorbed
BTU/HR
3,425,522
3,053,164
2,448,559
5,723,143
---
---
---
---
---
Relief Rate
LB/HR
28,357
25,275
20,270
47,377
---
---
---
---
---
LB/HR
121,278
TOTAL
TOTAL RELIEF RATE
Temperature
°F
173.6
Viscosity
cP
0.0000
Z
0.7724
MW
Compressibility
51
k=Cp/Cv Ratio
Lower Heating Value
1.2735
BTU/SCF
2,575
<===
Contract :
04411700
Revision :
A
Date :
23-Jun-2000
By:
Figure D1.14
Your Name Here
Chk:
Appv'd:
FIRE CASE RELIEF LOAD CALCULATIONS - VAPOR ONLY
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
CALCULATION METHOD
GENERAL
PROCESS:
PSV:
UNIT:
SET PRES:
EQUIP. NO.:
OVER PRES:
W = 0.1406 * (M * P1)^0.5 * (Tw - T1)^1.25 / T1^1.1506 * A
Where:
W = Relief Load, LB/HR
A = Exposed Surface Area of Vessel, Ft²
EQUIP. SVC.:
ATM PRES:
P&ID NO.:
RELIEF PRES:
M = Molecular Weight of Vapor
PSIG
21 %
14.70 PSI
14.70 PSIA
NOTES
P1 = Relieving Pressure, PSIA
1. Fire exposed vessel area is exposed area up to a maximum of 25 FT from grade.
Tw = Vessel Wall Temperature, °R
2. Maximum vessel wall temperature recommended as 1557 °R for carbon steel; adjust
T1 = Relieving Temperature, °R = Tn * P1 / Pn
for other materials to reflect maximum recommended temperatures. Default value is
Tn = Normal Operating Temperature, °R
1557 °R. If relieving temperature exceeds vessel wall temperature, then PSV will
Pn = Normal Operating Pressure, PSIA
not protect vessel during fire.
3. Vapor properties at normal conditions.
Reference: API RP-521
INPUT DATA
EQUIPMENT IN FIRE ZONE TAG NO.
EQUIPMENT SERVICE
P&ID NO. (EQUIPMENT LOCATION)
VESSEL Type(1=H;2=V;3=S)
Number of Identical Vessels
Diameter, O.D.
FT
H or L
FT
Vessel Elevation to Bottom
FT
H or D affected by fire
FT
(Note 1)
Number of Heads affected by fire
VAPOR
Normal Oper. Press.
PSIG
Normal Oper. Temp.
°F
Vessel Wall Maximum Temp.
°F
MW
Compressibility
(Note 2)
(Note 3)
Z
k=Cp/Cv Ratio
Lower Heating Value
BTU/SCF
Viscosity
cP
CALCULATED RESULTS
VESSEL Surface Area
FT²
---
---
---
---
---
---
---
---
---
Relieving Temperature
°F
---
---
---
---
---
---
---
---
---
Heat Absorbed
BTU/HR
---
---
---
---
---
---
---
---
---
Relief Rate
LB/H
---
---
---
---
---
---
---
---
---
TOTAL RELIEF RATE
LB/H
0
Temperature
°F
0
Viscosity
cP
0
MW
Compressibility
0
Z
k=Cp/Cv Ratio
Lower Heating Value
0
0
BTU/SCF
0
<===
Figure D1.15
Contract :
04411700
Revision :
A
Date :
15-Jun-2000
By:
Chk:
Your name here
Appv'd:
HEAT EXCHANGER TUBE RUPTURE
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-15.xls]FIG D1.15
GENERAL
PLANT:
NAC
PSV:
UNIT:
SET PRES:
EQUIP. NO.:
OVER PRES:
EQUIP. SVC.:
TOLUENE COLUMN REBOILER
P&ID NO.:
FLARE
C-SV1011 A/B
RELIEF VALVE
3.5 KG/CM²G
10 %
ATM PRES:
1.03 KG/CM²A
RELIEF PRES:
4.9 KG/CM²A
INPUT DATA (NOTE 1)
FLUID PHASE: 1=LIQUID, 2=VAPOR, OR 3=TWO PHASE (NOTE 2)
2
TUBE I.D., d
0.782 IN
DESIGN PRESSURE, HIGH PRESSURE SIDE
20 KG/CM²G
NORMAL OPERATING PRESSURE, HIGH PRESSURE SIDE, P1
15.8 KG/CM²G
NORMAL OPERATING TEMPERATURE, HIGH PRESSURE SIDE
203 °C
WEIGHT % FLASH, R (NOTE 3)
100 %
BUBBLE POINT PRESSURE, Pbp (NOTE 4)
15.8 KG/CM²A
LIQUID SG AT FLOWING CONDITIONS
1.000
VAPOR MW
18.02
VAPOR COMPRESSIBILITY, Z
HIGH PRESSURE
SIDE
HEAT EXCHANGER
1
VAPOR SPECIFIC HEAT RATIO, k
1.30
ORIFICE COEFFICIENT, C
LOW PRESSURE
SIDE
0.60
SINGLE PHASE RESULTS
API RP-521 PRESSURE RATIO TWO-THIRDS RULE:
L.P. Side @ 4.9 KG/CM²A < 2/3*H.P. Side @ 14.0 KG/CM²G
TUBE AREA
3.10 CM²
EXPANSION COEFFICIENT, Y
1. The following basis is used to determine the required relieving flow rate: the tube failure is a sharp break in one tube;
the tube failure occurs at the back side of the tube sheet; high pressure fluid flows both through the tube stub remaining
in the tube sheet and through the other, longer section of the tube. For simplification, the calculation equations are
CRITICAL HYDRAULIC PRESSURE, Pcf
9.2 KG/CM²A
VAPOR FLOW REGIME
CRITICAL
VAPOR DENSITY AT OPERATING CONDITIONS, HIGH PRESSURE SIDE, pv
2. Liquid flow is calculated based on incompressible flow assuming an orifice coefficient of 0.6; Vapor flow at critical
7.51 KG/M³
LIQUID DENSITY AT OPERATING CONDITIONS, HIGH PRESSURE SIDE, pv
TWO PHASE RESULTS
CRITICAL BUBBLE POINT HYDRAULIC PRESSURE
DELTA PRESSURE FOR TWO PHASE FLOW
LIQUID RELIEF FLOW RATE COEFFICIENT, CL
conditions is calculated using the API RP-520 equation with an orifice coefficient of 0.6 and a beta ratio of less than 0.2;
Vapor flow at sub-critical conditions uses a more generalized form of the equation; Two phase flow uses the
SINGLE PHASE LIQUID RELIEF FLOW RATE AT PSV INLET CONDITIONS, Wl
SINGLE PHASE VAPOR RELIEF FLOW RATE AT PSV INLET CONDITIONS, Wv
conservatively based on flow through two orifices.
methodology recommended in API RP-520 for sizing two-phase relief valves.
3,148 KG/HR
3. Flash the high pressure side fluid at the greater of the critical bubble point hydraulic pressure or the low pressure side
pressure to determine the weight percent of the fluid that is vapor. Alternatively, to simplify calculations,
a conservative solution can be reached by assuming no flashing occurs across the ruptured tube. Actual relief valve
sizing will be based on flashing at relief conditions.
4. If the high pressure side fluid is all liquid at normal conditions, but flashes as it passes through the orifice, calculate the
VAPOR RELIEF FLOW RATE COEFFICIENT, CV
bubble point pressure at the normal temperature of the high pressure side fluid. If the high pressure side fluid is two
AREA, LIQUID PHASE
phase at normal conditions, then enter the high pressure side normal operating pressure as the bubble point pressure.
AREA, VAPOR PHASE
TWO PHASE LIQUID RELIEF FLOW RATE
TWO PHASE VAPOR RELIEF FLOW RATE AT PSV INLET CONDITIONS
Contract :
04411700
Revision :
A
Date :
FIGURE D1.16
By:
23-Jun-2000
Your name here
Chk:
Appv'd:
HOU FLARE STACK: RAD. INT. AT WASTE WATER TREAT. FOR UNI. CRACK. DEP. RATE
FLARE RADIATION CALCULATIONS
ENTER ENGLISH UNITS INPUT DATA BELOW
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
Flare Gas Data
P
Absolute Pressure
Tg
Temperature
W
Flow Rate
k
NHV
CL
Mg
English Units
15 psia
Specific Heat Ratio = Cp/Cv
264.6 °F
129.2 °C
112,000 kg/hr
13,465 Btu/lb
Lean flammability Limit (vol % or mol %): enter fraction value
4.30%
1.30
7,481 kcal/kg
4.30%
Molecular Weight
5.70 lb/lbmol
5.70 g/gmol
F
Fraction of Heat Recieved as Radiant Heat
0.15
0.15
t
Fraction of Radiant Heat Transmitted
1.0
1.0
Flare Stack Dimensional Data
d
Flare Tip Diameter
66.0 in
1.7 m
H
Flare Stack Height
263.0 ft
80.2 m
h
Vertical Distance from Stack Base to Object (negative if object below stack base)
-27.2 ft
-8.3 m
R
Horizontal Distance from Stack Base to Object
154.2 ft
47.0 m
Ua
Wind Velocity
Ma
Air Molecular Weight
Ta
Air Temperature
Wind/Air Data
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
NOTES:
1. If NHV is not available, GO TO NHV SHEET FOR ESTIMATE VALUE.
29.3 ft/sec
28.96 lb/lbmol
70.0 °F
Note
1.05 kg/cm^2(A)
246,919 lb/hr
1.30
Net Heating Value
Metric Units
8.9 m/sec
28.96 g/gmol
21.1 °C
1
Contract :
04411700
Revision :
A
Date :
FIGURE D1.17
By:
23-Jun-2000
Your name here
Chk:
Appv'd:
#4 FLARE STACK (TOTAL POWER FAILURE CASE)
FLARE RADIATION CALCULATIONS
ENTER METRIC UNITS INPUT DATA BELOW
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
Flare Gas Data
P
Absolute Pressure
Tg
Temperature
W
Flow Rate
k
Specific Heat Ratio = Cp/Cv
NHV
CL
Mg
Metric Units
1.03 kg/cm^2(A)
129.2 °C
1,287,363 kg/hr
1.28
Net Heating Value
7,481 kcal/kg
Lean flammability Limit (vol % or mol %): enter fraction value
1.58%
Molecular Weight
65.50 g/gmol
F
Fraction of Heat Recieved as Radiant Heat
t
Fraction of Radiant Heat Transmitted
English Units
264.6 °F
2,838,120 lb/hr
1.28026331
13,465 Btu/lb
1.58%
65.50 lb/lbmol
0.15
0.15
1
1
Flare Stack Dimensional Data
d
Flare Tip Diameter
1.5 m
59.25 in
H
Flare Stack Height
121.5 m
398.6 ft
h
Vertical Distance from Stack Base to Object (negative if object below stack base)
-28.3 m
-92.8 ft
R
Horizontal Distance from Stack Base to Object
133.2 m
437.0 ft
Ua
Wind Velocity
Ma
Air Molecular Weight
Ta
Air Temperature
Wind/Air Data
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
NOTES:
1. If NHV is not available, GO TO NHV SHEET FOR ESTIMATE VALUE.
8.9 m/sec
28.96 g/gmol
21.1 °C
Note
14.70 psia
29.2 ft/sec
28.96 lb/lbmol
70.0 °F
1
FIGURE D1.18
Contract :
04411700
Revision :
A
X_l
Date :
23-Jun-2000
By:
FLARE RADIATION CALCULATIONS (B&S METHOD)
Your name here
Chk:
Appv'd:
WIND
L
S_l
#4 FLARE STACK (TOTAL POWER FAILURE CASE)
NOTES:
θ
Z_l
1. Brzustowski's and Sommer method for flare radiation calculation is taken from API Recommended,
Practice, 3rd ed., Nov. 1990.
2. Ideal Gas Law is used.
Xm - H
3. Lean flammability limit is taken from pure components values adjusted for the mol. wt. of flared gas.
Yc
XXcc
RADIATION INTENSITY (K) VERSUS HORIZONTAL DISTANCE FROM STACK BASE TO OBJECT (R)
K==========> R Table
d
R==========> K Table
K
R
K
R
R
K
R
K
kcal/hr/m
m
Btu/hr/ft
ft
m
kcal/hr/m
ft
Btu/hr/ft
1,400
241
2,700
4,100
500
79
813
1,000
#NUM!
1,500
0
250
#NUM!
2,707
0
998
50
2,797
100
1,033
65
2,759
200
1,022
8,200
#NUM!
3,000
#NUM!
100
2,569
213
1,017
13,600
#NUM!
5,000
#NUM!
150
2,153
300
968
Rmin = Xc
Kmax
m
kcal/hr/m
39
PEAK RADIATION INTENSITY
2,806
Rmin = Xc
Kmax
ft
Btu/hr/ft
127
Flare Gas Data
P
Absolute Pressure
Tg
Temperature
W
Flow Rate
k
NHV
D
H
H' = H + Yc
N = R - Xc
1,034
R
Metric Units
1.03 kg/cm^2(A)
Specific Heat Ratio = Cp/Cv
129.20 °C
264.56 °F
2,838,120 lb/hr
7,481 kcal/kg
CL
Lean flammability Limit (vol % or mol %): enter fraction value
Mg
Molecular Weight
0.02
65.50 g/gmol
Note
14.70 psia
1,287,363 kg/hr
1.28
Net Heating Value
English Units
1.28
13,465 Btu/lb
0.02
3
65.50 lb/lbmol
F
Fraction of Heat Recieved as Radiant Heat
0.15
0.15
t
Fraction of Radiant Heat Transmitted
1.00
1.00
Flare Stack Dimensional Data
d
Flare Tip Diameter
1.5 m
59.25 in
H
Flare Stack Height
121.5 m
398.62 ft
h
Vertical Distance from Stack Base to Object (negative if object below stack base)
-28.3 m
-92.85 ft
R
Horizontal Distance from Stack Base to Object
133.2 m
437.01 ft
Ua
Wind Velocity
Ma
Air Molecular Weight
Ta
Air Temperature
21.1 °C
ρ
Air Density (kg/m3) = Ma*P*/(0.084784*(Ta + 273.16)).
1.20 kg/m
0.07 lb/ft
2
1.98 kg/m
0.12 lb/ft
2
Wind/Air Data
8.9 m/sec
28.96
29.20 ft/sec
28.96
69.98 °F
3
Air Density (lb/ft ) = Ma*P/(10.732*(Ta + 459.67))
Flare Gas Calculations
ρ
Gas Density (kg/m3) = Mg*P*/(0.084784*(Tg + 273.16))
Gas Density (lb/ft3) = Mg*P*/(10.732*(Tg + 459.67))
V
Volumetric Gas Flow Rate = W/ gas density/3600
180 mü/sec
6,366 ftü/sec
Ug
Gas Exit Velocity = 4V/(pi*d^2) [m/sec] or 4*144V/(pi*d^2) [ft/sec]
101 m/sec
332 ft/sec
Us
Gas Sonic Velocity = 91.2(k*(Tg+273.16)/Mg)^0.5 [m/sec] or 223(k*(Tg+459.67)/Mg)^0.5 [ft/sec]
256 m/sec
839 ft/sec
Ug/Us
0.40
Mach #
dP
Pressure Drop at Flare Tip (m H2O) = (8.27*10^-7*W*((Tg+273.16)/Mg)^0.5/d^2)^2
1.4 m H2O
0.40
53.5 in H2O
Pressure Drop at Flare Tip (in H2O)= (2.72*10^-3*W*((Tg+459.67)/Mg)^0.5/d^2)^2
Q
Heat Released = W * NHV
9.63E+09 kcal/hr
3.82E+10 Btu/hr
Qr
Heat Transmitted = Q * F (emissivity)
1.44E+09 kcal/hr
5.73E+09 Btu/hr
Dimensionaless Parameters Calculations
C_L
CL*(Ug/Ua)*(Mg/Ma)
40.7%
40.7%
S_L
If C_L <=0.5, S_L = 2.04/(C_L^1.03). If C_L > 0.5, S_L = 2.51/(C_L^0.625)
5.15
5.15
X_L
If C_L <=0.5, X_L = S_L - 1.65. If C_L > 0.5 and S_L > 2.35, X_L = S_L - 1.65.
3.50
3.50
If C_L > 0.5 and S_L <= 2.35, X_L = 10^((log(S_L) - log (1.04) - log (2.05))/2.28)
Z_L
2.05*X_L^0.28
R_
(Ug/Ua)*(Gas density/Air density)^0.5
2.91
2.91
14.64
14.64
Z_l
Z_L * d * R_L [m] or Z_L* d * R_L/12 [ft]
64.2 m
X_l
X_L * d * R_L [m] or X_L * d * R_L/12 [ft]
77.2 m
253.1 ft
Xc
X_l/2
38.6 m
126.6 ft
Yc
Z_I * 0.82
52.6 m
172.6 ft
Dimensional Parameters Calculations
210.5 ft
Flame Parameters Calculations
q
Angle of Flare Flame from Vertical = ATAN (Xc/Yc)
L
Flame Length = SQRT(Z_l^2 + X_l^2)
100.3 m
329.2 ft
Actual Curved Lengh of Flame = S_L * d * R_L [m] or S_L * d * R_L/12 [ft]
113.5 m
372.4 ft
S_l
36.2
36.2
Single Point Radiation Parameters Calculations
N
Horizontal Distance from Flame Radiation Source Point to Object = |R - Xc|
H'
Vertical Distance from Object to the Flame Radiation Source Point = |H + Yc - h|
D
Diagonal Distance from Flame Radiation Source Point to Object = (N^2 + H'^2)^0.5
K
Radiation Intensity = (Tau * Qr)/(4pi *D^2)
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
94.6 m
310.4 ft
202.4 m
664.1 ft
223.4 m
733.1 ft
2,303 kcal/hr/m
849 Btu/hr/ft
FIGURE D1.19
Contract :
04411700
Revision :
A
X_l
Date :
23-Jun-2000
By:
FLARE RADIATION CALCULATIONS (SIMPLE METHOD)
Your name here
Chk:
Appv'd:
L
WIND
S_l
#4 FLARE STACK (TOTAL POWER FAILURE CASE)
NOTES:
θ
Z_l
1. Simple method for flare radiation calculation is taken from FDI Radiation Calculation Spreadsheet developed by
Ed Aitken and API Recommended Practice 521, 3rd ed., Nov. 1990.
2. Lean flammability is not required to perform radiation calculations.
Xm - H
3. Ideal Gas Law is used.
Yc
4. Flame length equation is derived from flame length vs. Q curve. The difference between the equation and curve is
about plus or minus 3%.
XXcc
RADIATION INTENSITY (K) VERSUS HORIZONTAL DISTANCE FROM STACK BASE TO OBJECT (R)
K==========> R Table
d
R==========> K Table
K
R
K
R
R
K
R
K
kcal/hr/m
m
Btu/hr/ft
ft
m
kcal/hr/m
ft
Btu/hr/ft
1,400
281
2,700
4,100
500
138
1,000
#NUM!
1,500
941
0
447
#NUM!
2,670
0
984
50
2,987
100
1,068
65
3,018
200
1,111
8,200
#NUM!
3,000
#NUM!
100
2,954
213
1,112
13,600
#NUM!
5,000
#NUM!
150
2,592
300
1,101
Rmin = Xc
Kmax
m
kcal/hr/m
71
Rmin = Xc
Kmax
ft
Btu/hr/ft
PEAK RADIATION INTENSITY
3,020
232
D
H
H' = H + Yc
N = R - Xc
1,113
R
Flare Gas Data
P
Absolute Pressure
Tg
Temperature
W
Flow Rate
k
Specific Heat Ratio = Cp/Cv
NHV
Metric Units
1.03 kg/cm^2(A)
129.20 °C
264.56 °F
2,838,120 lb/hr
1.28
7,481 kcal/kg
CL
Lean flammability Limit (vol % or mol %)
Mg
Molecular Weight
Note
14.70 psia
1,287,363 kg/hr
1.28
Net Heating Value
English Units
0.02
13,465 Btu/lb
0.02
65.50 g/gmol
2
65.50 lb/lbmol
F
Fraction of Heat Recieved as Radiant Heat
0.15
0.15
t
Fraction of Radiant Heat Transmitted
1.00
1.00
Flare Stack Dimensional Data
d
Flare Tip Diameter
1.5 m
59.3 in
H
Flare Stack Height
121.5 m
398.6 ft
h
Vertical Distance from Stack Base to Object (negative if object below stack base)
-28.3 m
-92.8 ft
R
Horizontal Distance from Stack Base to Object
133.2 m
437.0 ft
Ua
Wind Velocity
Ma
Air Molecular Weight
29.0
29.0
Ta
Air Temperature
21.1 °C
70.0 °F
ρ
Air Density (kg/m3) = Ma*P*/(0.084784*(Ta + 273.16)).
Air Density (lb/ft3) = Ma*P/(10.732*(Ta + 459.67))
1.20 kg/m
0.07 lb/ft
3
ρ
Gas Density (kg/m3) = Mg*P*/(0.084784*(Tg + 273.16))
1.98 kg/m
0.12 lb/ft
3
Wind/Air Data
8.9 m/sec
29.2 ft/sec
Flare Gas Calculations
Gas Density (lb/ft3) = Mg*P*/(10.732*(Tg + 459.67))
V
Volumetric Gas Flow Rate = W/ gas density/3600
180 mü/sec
6,366 ftü/sec
Ug
Gas Exit Velocity = 4V/(pi*d^2) [m/sec] or 4*144V/(pi*d^2) [ft/sec]
101 m/sec
332 ft/sec
Us
Gas Sonic Velocity = 91.2(k*(Tg+273.16)/Mg)^0.5 [m/sec] or 223(k*(Tg+459.67)/Mg)^0.5 [ft/sec]
256 m/sec
839 ft/sec
Ug/Us
0.40
0.40
Mach #
dP
Pressure Drop at Flare Tip (m H2O) = (8.27*10^-7*W*((Tg+273.16)/Mg)^0.5/d^2)^2
1.4 m H2O
53.5 in H2O
Pressure Drop at Flare Tip (in H2O)= (2.72*10^-3*W*((Tg+459.67)/Mg)^0.5/d^2)^2
Q
Heat Released = W * NHV
9.63E+09 kcal/hr
3.82E+10 Btu/hr
Qr
Heat Transmitted = Q * F (emissivity)
1.44E+09 kcal/hr
5.73E+09 Btu/hr
Dimensionaless Parameters Calculations
Ua/Ug
Z/L
Wind Velocity to Gas Exit Velocity Ratio
If Ua/Ug > 0.45, Z/L = 0.09. If Ua/Ug <= 0.05, Z/L = -8*Ua/Ug+1.
0.088
0.088
0.49
0.49
0.77
0.77
90.6
297.2 ft
If Ua/Ug > 0.05 and Ua/Ug <= 0.45, Z/L = 0.093*(Ua/Ug)^-0.684
X/L
If Ua/Ug > 0.45, X/L = 0.97. If Ua/Ug <= 0.05, X/L = 12*Ua/Ug.
If Ua/Ug > 0.05 and Ua/Ug <= 0.45, X/L = 1.082*(Ua/Ug)^0.142
Dimensional Parameters Calculations
Z_l
Z/L * L
X_l
X/L * L
Xc
X_l/2
Yc
Z_I/2
141.3 m
463.7 ft
70.7 m
231.8 ft
45.3 m
148.6 ft
Flame Parameters Calculations
q
Angle of Flare Flame from Vertical = ATAN (Xc/Yc)
L
Flame Length = 0.5667 * 10^(0.45 log Q - 1.98) [m] or 10^(0.45 log Q - 1.98) [ft]
m
57.3
57.3
184.5
605.4 ft
Single Point Radiation Parameters Calculations
m
N
Horizontal Distance from Flame Radiation Source Point to Object = |R - Xc|
H'
Vertical Distance from Object to the Flame Radiation Source Point = |H + Yc - h|
195.1
640.1 ft
D
Diagonal Distance from Flame Radiation Source Point to Object = (N^2 + H'^2)^0.5
204.9 m
672.1 ft
K
Radiation Intensity = (Tau * Qr)/(4pi *D^2)
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
62.5 m
2,739 m
205.2 ft
1,010 Btu/hr/ft
4
FIGURE D1.20
Contract :
04411700
Revision :
A
X_l
Date :
23-Jun-2000
By:
FLARE RADIATION CALCULATIONS (KENT METHOD)
Your name here
Chk:
Appv'd:
WIND
L
S_l
#4 FLARE STACK (TOTAL POWER FAILURE CASE)
NOTES:
θ
Z_l
1. Kent method for flare radiation calculation is taken from Flare Gas Systems Pocket Handbook by
K. Banerjee, N. P. Cheremisinoff, and P. N. Cheremisinoff, 1985.
2. Lean flammability is not required to perform radiation calculations.
Xm - H
Yc
XXcc
RADIATION INTENSITY (K) VERSUS HORIZONTAL DISTANCE FROM STACK BASE TO OBJECT (R)
K==========> R Table
K
R
kcal/hr/m
m
1,400
K
R
Btu/hr/ft
191
d
R==========> K Table
R
ft
500
m
650
K
R
kcal/hr/m
ft
0
0
886
2,700
#NUM!
1,000
#NUM!
2,311
100
875
4,100
#NUM!
1,500
#NUM!
65
2,241
200
834
8,200
#NUM!
3,000
#NUM!
100
2,030
213
826
13,600
#NUM!
5,000
#NUM!
150
1,677
300
769
Rmin = Xc
Kmax
m
kcal/hr/m
6
50
K
Rmin = Xc
Kmax
ft
Btu/hr/ft
PEAK RADIATION INTENSITY
2,404
20
D
Btu/hr/ft
2,402
H
H' = H + Yc
N = R - Xc
886
R
Flare Gas Data
P
Absolute Pressure
Tg
Temperature
W
Flow Rate
k
Specific Heat Ratio = Cp/Cv
NHV
Metric Units
1.03 kg/cmý(A)
129.20 °C
264.56 °F
2,838,120 lb/hr
7,481 kcal/kg
Lean flammability Limit (vol % or mol %)
Mg
Molecular Weight
0.02
65.50 g/gmol
1.28
13,465 Btu/lb
0.02
65.50 lb/lbmol
F
Fraction of Heat Recieved as Radiant Heat
0.15
0.15
t
Fraction of Radiant Heat Transmitted
1.00
1.00
Flare Stack Dimensional Data
d
Flare Tip Diameter
1.5 m
59.3 in
H
Flare Stack Height
121.5 m
398.6 ft
h
Vertical Distance from Stack Base to Object (negative if object below stack base)
-28.3 m
-92.8 ft
133.2 m
437.0 ft
R
Horizontal Distance from Stack Base to Object
Ua
Wind Velocity
Ma
Air Molecular Weight
29.0
Ta
Air Temperature
21.1 °C
70.0 °F
1.20 kg/m
0.07 lb/ft
Wind/Air Data
8.9 m/sec
ρ
Air Density (kg/m ) = Ma*P*/(0.084784*(Ta + 273.16)).
Air Density (lb/ft3) = Ma*P/(10.732*(Ta + 459.67))
ρ
Gas Density (kg/m3) = Mg*P*/(0.084784*(Tg + 273.16))
3
29.2 ft/sec
29.0
Flare Gas Calculations
1.98 kg/m
0.12 lb/ft
3
Gas Density (lb/ft ) = Mg*P*/(10.732*(Tg + 459.67))
V
Volumetric Gas Flow Rate = W/ gas density/3600
180 mü/sec
6,366 ftü/sec
Ug
Gas Exit Velocity = 4V/(pi*d^2) [m/sec] or 4*144V/(pi*d^2) [ft/sec]
101 m/sec
332 ft/sec
Us
Gas Sonic Velocity = 91.2(k*(Tg+273.16)/Mg)^0.5 [m/sec] or 223(k*(Tg+459.67)/Mg)^0.5 [ft/sec]
256 m/sec
839 ft/sec
Ug/Us
0.40
0.40
Mach #
dP
Pressure Drop at Flare Tip (m H2O) = (8.27*10^-7*W*((Tg+273.16)/Mg)^0.5/d^2)^2
1.4 m H2O
53.5 in H2O
Pressure Drop at Flare Tip (in H2O)= (2.72*10^-3*W*((Tg+459.67)/Mg)^0.5/d^2)^2
Q
Heat Released = W * NHV
Qr
Heat Transmitted = Q * F (emissivity)
9.63E+09 kcal/hr
3.82E+10 Btu/hr
1.44E+09 kcal/hr
5.73E+09 Btu/hr
Dimensionaless Parameters Calculations
Ua/Ug
Wind Velocity to Gas Exit Velocity Ratio
0.088
0.088
190.6 m
625.4 ft
Dimensional Parameters Calculations
Xm
[H*(H+L)]^0.5
Xc
(Xm - H)sin(flame angle)
6.0 m
19.8 ft
Yc
(Xm - H)cos(flame angle)
68.9 m
225.9 ft
q
Angle of Flare Flame from Vertical = ATAN (Ua/Ug)
L
Flame Length = 118*d [m] or 118*d/12 [ft]
Flame Parameters Calculations
5.0
5.0
177.6 m
582.6 ft
Single Point Radiation Parameters Calculations
N
Horizontal Distance from Flame Radiation Source Point to Object = |R - Xc|
127.2 m
417.2 ft
H'
Vertical Distance from Object to the Flame Radiation Source Point = |H + Yc - h|
218.7 m
717.4 ft
D
Diagonal Distance from Flame Radiation Source Point to Object = (N^2 + H'^2)^0.5
252.9 m
829.9 ft
K
Radiation Intensity = (Tau * Qr)/(4pi *D^2)
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-16-21.xls]FIG D1.20-KENT Method
1,797 kcal/hr/m
Note
14.70 psia
1,287,363 kg/hr
1.28
Net Heating Value
CL
English Units
662 Btu/hr/ft
2
FIGURE D-21
LHV OR NHV CALCULATION BASED ON AVERAGE MOLECULAR WEIGHT:
BTU/SCF = 50(M.W.) + 100
BTU/LB = (50(M.W.) + 100)*379/M.W.
M.W. = molecular weight
M.W. =
8.70
BTU/SCF
535
BTU/LB
23,306
KCAL/KG
12,948
---------------- ---------------- ---------------- ---------------- ---------------#4 FLARE STACK (TOTAL POWER FAILURE CASE)
LHV OR NHV CALCULATION BASED ON COMPOSITIONS:
#DIV/0!
PLANT
PSV TAG NO.EQUIPMENT
KG/HR
#3 MDU
SV5021
FUEL GAS K.O. DRUM
1,007
PSV1515A/B CATALYTIC COLUMN
13,275
MTBE
MTBE
PSV1524
MeOH/WATER COLUMN
2,273
#4 MDU
PSV6018
STRIPPER
13,722
1,066
#3 SRP
PSVC3464 NO. 4 DEA ABSORBER
#3 PLAT
PSV2507
STABILIZER
99,997
#3 PLAT
PSV2516
FUEL GAS K.O. DRUM
5,248
XYLENE FRACT
PSV2806
DEPENTANIZER 43,532
#5 GAS CONC
PSV2701
DEETHANIZER 11,221
#5 GAS CONC
PSV2702
C3/C4 SPLITTER 14,744
#3 GAS CONC
PSV2762
DEETHANIZER 16,284
#3 GAS CONC
PSV2765
C3/C4 SPLITTER 33,664
#4 CDU
SV4001A-F CRUDE COLUMN467,214
#4 CDU
SV4008A/B SPLITTER
128,174
#4 CDU
SV4010
STABILIZER
140,820
FUEL GAS BALANCE
6,103
DRUM
#4 CDU
SV4018
#4 GAS CONC
SV4201
DEETHANIZER 47,248
#4 GAS CONC
SV4203
DEPROPANIZER 73,327
BRU
168,000
#2 MDU
PSV2001
HP SEPARATOR
444
M.W.
NHV (kcal/NMü)
NHV (kcal/kg)k (Cp/Cv)
13,465
7,481
9,186
25,344
7,201
3,494
14,990
25,185
2,155
34,751
19,304
21,641
18,664
21,777
10,199
27,284
26,347
34,782
18,126
21,689
45,592
5,029
12,321
10,436
5,041
3,143
11,306
10,954
2,108
10,781
11,145
11,068
11,002
11,061
2,315
9,122
10,929
10,806
10,974
11,067
10,750
12,948
1.230
1.073
1.185
1.223
1.181
1.076
1.350
1.060
2.015
1.450
2.254
1.434
1.062
1.113
1.210
1.128
2.401
1.424
1.100
1.310
# mol
mol %
0.1
0.0
0.3
1.0
0.0
1.2
0.2
0.0
0.4
1.1
0.0
2.8
0.1
0.0
0.2
7.8
0.2
9.9
0.4
0.0
1.2
3.4
0.0
3.1
0.9
0.0
1.5
1.1
0.0
1.7
1.3
0.0
2.2
2.6
0.1
3.9
36.3
0.4
24.1
10.0
0.1
9.7
10.9
0.2
13.3
0.5
0.0
0.4
3.7
0.1
6.5
5.7
0.1
8.5
13.0
0.1
9.0
0.0
0.0
0.3
0.0
#DIV/0!
#DIV/0!
0.0
#DIV/0!
#DIV/0!
---------------- ---------------- ---------------- ---------------- ---------------- ---------------- ---------------- ---------------- ---------------- ---------------- ---------------TOTAL
1,287,363
7,481
1.280
100.0
1.5
100.0
BTU/LB
KCAL/KG
16.7
54.4
32.0
24.9
29.7
51.5
22.9
72.2
38.8
43.8
38.0
44.1
98.7
67.0
54.0
72.1
37.0
43.9
95.0
8.7
wt%
k_i
0.0037726
0.01331858
0.00428144
0.03428207
0.00215612
0.10627094
0.01573673
0.03250871
0.02964131
0.02482743
0.04913081
0.05567994
0.25570883
0.10830358
0.16050127
0.00485668
0.15595395
0.1209851
0.0989466
0.00340062
#DIV/0!
#DIV/0!
---------------1.3
Figure D1.22
Contract :
04411700
Revision :
A
Date :
15-Jun-2000
By:
Your name here
Chk:
Appv'd:
HORIZONTAL K.O. RATING SPREADSHEET
NRP FLARE K.O. DRUM (Y-V2907)
TOTAL POWER FAILURE CASE
DRUM DATA
Length
14.7 m
48.2 ft
4.9 m
16.1 ft
Radius
2.45 m
8.04 ft
Liquid Level
1.23 m
4.0 ft
Vapor Space
3.7 m
12.1 ft
Diameter
FLUID DATA
Temperature
Pressure
M.W.
Liquid Density
Gas Viscosity
Gas Density
Liq. Drop. Vel.
Liq. Drop. Time
Vap. Vel.
77.1
2.50
48.6
700
0.01
4.10
0.3
13.6
1.1
øC
kg/cmýA
kg/m
cp
kg/m
m/s
s
m/s
170.8
35.68
48.6
43.7
0.01
0.26
0.9
13.5
3.6
øF
psia
lb/ft
cp
lb/ft
ft/s
s
ft/s
PROBLEM TYPE
Capacity Known
Capacity
Particle Size
YES enter capacity (kg/hr) below
240,600 kg/hr
115 microns
Next Guess
Converge
converge
530,427 lb/hr
0.00038 ft
114
PARAMETERS CALCULATIONS
C(Re)
5.6E+02
Drag Coeff. (C)
3.44
Tot. Flow Area
18.9 m
Area for Liq. Flow
3.7 m
Area for Vap. Flow
15.2 m
Liq. Volume
54 m
% capacity diff.
1.0
RESULTS SUMMARY
Drum Capacity
243,130 kg/hr
Particle Size
115 microns
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-22.xls]FIG D1.22
5.7E+02
3.43
203.0
39.7
163.3
1914
ft
ft
ft
ft
537,194 lb/hr
0.000377297 ft
Figure D1.23
USERS GUIDE:
PRESSURE RELIEF VALVE DATA SHEET
OVERVIEW
This spreadsheet is a data sheet to transfer relief load and other process information to control systems to
select suitable relief valve(s) for the specified service. This spreadsheet also includes a capability to
provide preliminary sizing calculations for relief valves and provide rated flow for checking inlet and
outlet hydraulics as recommended by API. This electronic file contains the following sheets:
1. Data Sheet: This is the primary sheet for transfer of information to control systems and is typically
given by control systems to the vendor for sizing. This sheet is also used for preliminary sizing
estimations & checks. A menu allows the user to choose the units, liquid PSV sizing method, and
whether estimated sizes are hidden or revealed. The menu button is invisible on printed copies.
2. Users Guide: Explanation for first time users.
3. Calculation Sheet \ Dialog Sheets \ Macro Sheet: These are primarily there to do the calculations.
They are laid out in a manner that can be followed, checked, updated, or changed; but usually this is
unnecessary.
Menu
STEP-BY-STEP WALKTHROUGH
1. Read users guide (good job!).
2. Find the “menu” button on the upper left of the Data Sheet, or to the right.
3. Choose English or Field Metric units using the menu. You can always enter in non-standard units
manually, but remember that this will not change any sizing calculations.
4. Choose the liquid sizing method. Typically, Non-ASME gives a more conservative value, but
ASME is often preferred for liquid trim valves (see Sizing, below).
5. Choose to reveal sizing. This will allow an estimation of the size to be determined for P&ID’s or
preliminary check rating.
6. Fill in the data sheet. In addition to the process conditions at each relief scenario, make sure to fill in
the Revision Log, Equipment Conditions, valve type and discharge location. Add any sketches or
notes in the blank box provided.
7. Print the data sheet with sizing on it. Then go back to the menu, hide sizing, and print out the data
sheet without sizing. Both should be self-checked and checked by the checker.
PRELIMINARY RELIEF VALVE SIZING
Sizing is according to API 520 Part 1, sections 4.3 to 4.7. These equations do not apply to valves with a
backpressure greater than 50%, and are not recommended for sizing pilot valves. Sizing equations are for
preliminary check sizing only. Vendor should confirm all sizes. All sizing is done in the Calculation
Sheet and is clearly identified. Refer to that sheet for questions on calculation equations.
The backpressure sizing factor used by the data sheet is based on a curve from API. Since the curves only
allow a backpressure up to 50%, extrapolation was used to estimate the curve from 50% to 60%, and
beyond 60% the sizing factor is not estimated. The curve is conservative. The largest discrepancy found in
testing was a group of bellows valves where one vendor allowed a Kb = 1.0 and API had a Kb = 0.68,
causing API to predict 47% more required area than that vendor. Some vendor curves have changed in
recent years, and a vendor catalog can be used to obtain a less conservative sizing factor.
The liquid sizing method can also have a huge impact on required areas (one PSV on the Irving Oil Project
had a 75% greater calculated required area with Non-ASME than with ASME). The ASME equations will
typically predict a lower required area, but are based upon water and air at standard conditions. Since
vendors typically use a modified version of one of these equations, it is difficult to predict the size when
the answers given by the two equations differ greatly. In general, the API equation is recommended for
sizing existing non-liquid trim valves, and the ASME equation for sizing new valves where liquid relief is
controlling and a liquid valve may be specified. The mathematical difference is a factor to account for the
slower popping action for non-liquid trim valves in liquid service.
Figure D1.24
REVISION LOG
PRESSURE RELIEF VALVE DATA SHEET &
PRELIMINARY SIZE CALCULATIONS
FLUID PHYSICAL PROPERTIES
AT RELIEF CONDITIONS
CLIENT:
PLANT:
SITE:
CONTRACT #:
IRVING OIL
REVISION
RFCC
0
SITE
CONTRACT NUMBER
DATE
ISSUE STATUS
Failure
Failure
Cooling
Fire
Water Failure
NOTES (NOTE 1)
Blocked
Reflux
Pump Around
Air Fan
Abnormal
Control
Instrument
Tube
Discharge
Failure
Failure
Loss
Heat Input
Valve Failure
Air Failure
Rupture
Other
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
VAPOR FLOWRATE, LB/HR
0
684,032
72,562
121,278
CHK'D
APP'D
GENERAL
RELIEF CONDITIONS
Total Power Partial Power
BY
BASIC DESIGN
571,463
TAG NUMBER:
PSV-00951/00952/00954
UNIT:
RFCC
NAME OF FLUID:
HC
PROTECTED EQUIP NO. (NOTE 6): C-42517
LIQUID FLOWRATE, GPM
PROTECTED EQUIP SERVICE:
DEBUTANIZER
P& ID/FRAME NO:
AX-D-425-123
WT % FLASH
ALLOWABLE OVERPRESS, % (NOTE 5)
16
RELIEF PRESSURE, PSIA
199.70
16
21
16
229.30
229.30
238.55
229.30
178.4
178.4
173.6
173.6
MOL WT VAPOR
50.50
50.50
51.00
51.00
COMPRESSIBILITY, Z
0.772
0.772
0.772
0.772
NORMAL:
167.00 PSIG @
382.0 °F
SPECIFIC HEAT RATIO, K
1.274
1.274
1.274
1.274
DESIGN:
185.00 PSIG @
°F
90.00
90.00
90.00
90.00
90.00
MAWP:
PSIG @
°F
1.00
1.00
1.00
1.00
1.00
27.75
2.94
4.69
22.98
RELIEF TEMPERATURE, °F
FLANGE RATING & FACING:
PIPE SPEC:
LIQUID S.G.
LIQUID VISCOSITY, CP
SET PRESSURE (NOTE 3):
TOTAL BACK PRESSURE, PSIG
BACK PRES SIZING FACTOR, Kb or Kw
ESTIMATED REQUIRED AREA, IN²
---
PSIG
BACK PRESSURE (NOTE 8)
na
WARNING MESSAGES
GENERAL NOTES:
185
EQUIPMENT CONDITIONS (NOTE 2)
SUPERIMPOSED:
PSIG
BUILT-UP:
REMARKS & SKETCHES
PSI
TOTAL:
90.0 PSIG
INLET PRESSURE
1. All conditions to be evaluated for each relieving case. If condition
PIPE LINE SIZE:
IN
does not apply mark as "N/A". If the load is considered zero or
CALCULATED PRESSURE LOSS:
PSI
significantly lower than other major loads mark as "NEG".
ALLOWABLE PRESSURE LOSS:
2. If more than one vessel is involved in a single relief system, use
lowest design pressure and maximum allowable working
SIZING RELIEF COND:
pressure (MAWP).
OVER PRESSURE USED:
3. Set pressure is the pressure at which the relief valve starts to open,
(14.7 PSIA) at sea level.
5. Allowable overpressure is: 3% for fired steam boilers; 16% for
multiple valves; 21% for fire conditions; 10% for all other conditions.
6. Principal equipment where the PSV is located. Other equipment
to be protected is shown in the accompanying sketch.
7. The valve capacity at 10% overpressure which is the rated capacity
for API inlet and outlet hydraulic checks.
8. If the total back pressure is not entered, the program assumes a
default backpressure sizing factor, Kb = 1.0.
9. Sizing equations are for preliminary sizing only. All sizes to be
WARNING MESSAGES
16 %
27.75 IN²
DISCHARGE TO:
VALVE TYPE:
CONVENTIONAL
HEADER
BALANCED
ATM.
X
PILOT
PSV SIZE:
API ORIFICE AREA:
RATED FLOW
X
PROCESS
8 T 10
26.00 IN²
609,898 LB/HR VAPOR
RATED FLOW IS CALCULATED AT 10% OVERPRESSURE
AND NO BACKPRESSURE FOR HYDRAULIC CHECKS
PSV SIZE:
T
API ORIFICE AREA:
26.00 IN²
RATED FLOW
609,898 LB/HR VAPOR
MULTIPLE PSV INFORMATION
confirmed by vendor.
10. Preliminary liquid sizing method is AMSE.
PARTIAL POWER FAILURE
TOTAL API ORIFICE AREA REQ'D:
usually design pressure of equipment in PSIG.
4. Relief pressure is the set pressure + allowable overpressure +
5.55 PSI
PRELIMINARY RELIEF VALVE SIZING
PSV SIZE OPERATING
SPARE
AREA
8 T 10
IN²
T
IN²
A. Set pressure is less than 50 PSIG, consult a vendor.
TOTAL API ORIFICE AREA AVAIL.:
IN²
B. BALANCED BELLOWS PSVs: Backpressure exceeds 50% of the
TOTAL API ORIFICE AREA REQ'D:
IN²
set pressure, consult a vendor. The back pressure sizing factor
Kb is based upon API or an extrapolation of API.
C. CONVENTIONAL PSVs: Back pressure exceeds 10% of the set
pressure, consult a vendor. The backpressure sizing factor Kb is
1.0 for back pressures up to 10% of the set pressure.
Notice: This document has not been published and is the sole property of the Fluor Corporation, and is lent to the borrower for his confidential use only: and in consideration of the loan of this drawing, the borrower promises and agrees to return it upon request and agrees that it shall not be reproduced, copied, lent or otherwise disposed of directly or indirectly, nor used
for any purpose other than that for which it is specifically furnished.
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-23-24.xls]FIG D1.24-Data Sheet
FORM E-516i (REV 9/30)
Figure D1.25
REVISION LOG
PRESSURE RELIEF VALVE DATA SHEET
FLUID PHYSICAL PROPERTIES
AT RELIEF CONDITIONS
CLIENT:
PLANT:
SITE:
CONTRACT #:
IRVING OIL
REVISION
RFCC
0
SITE
CONTRACT NUMBER
ISSUE STATUS
DATE
Failure
Failure
Cooling
Fire
Water Failure
NOTES (NOTE 1)
Blocked
Reflux
Pump Around
Air Fan
Abnormal
Control
Instrument
Tube
Discharge
Failure
Failure
Loss
Heat Input
Valve Failure
Air Failure
Rupture
Other
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
VAPOR FLOWRATE, LB/HR
0
684,032
72,562
121,278
CHK'D
APP'D
GENERAL
RELIEF CONDITIONS
Total Power Partial Power
BY
BASIC DESIGN
571,463
TAG NUMBER:
PSV-00951/00952/00954
UNIT:
RFCC
NAME OF FLUID:
HC
PROTECTED EQUIP NO. (NOTE 6): C-42517
LIQUID FLOWRATE, GPM
PROTECTED EQUIP SERVICE:
DEBUTANIZER
P& ID/FRAME NO:
AX-D-425-123
WT % FLASH
ALLOWABLE OVERPRESS, % (NOTE 5)
RELIEF PRESSURE, PSIA
16
199.70
16
21
16
229.30
229.30
238.55
229.30
178.4
178.4
173.6
173.6
MOL WT VAPOR
50.50
50.50
51.00
51.00
COMPRESSIBILITY, Z
0.772
0.772
0.772
0.772
NORMAL:
167.00 PSIG @
382.0 °F
SPECIFIC HEAT RATIO, K
1.274
1.274
1.274
1.274
DESIGN:
185.00 PSIG @
°F
PSIG @
°F
RELIEF TEMPERATURE, °F
FLANGE RATING & FACING:
PIPE SPEC:
LIQUID S.G.
LIQUID VISCOSITY, CP
SET PRESSURE (NOTE 3):
185
PSIG
EQUIPMENT CONDITIONS (NOTE 2)
MAWP:
BACK PRESSURE
SUPERIMPOSED:
PSIG
BUILT-UP:
GENERAL NOTES:
REMARKS & SKETCHES
PSI
TOTAL:
90.0 PSIG
INLET PRESSURE
1. All conditions to be evaluated for each relieving case. If condition
PIPE LINE SIZE:
IN
does not apply mark as "N/A". If the load is considered zero or
CALCULATED PRESSURE LOSS:
PSI
significantly lower than other major loads mark as "NEG".
ALLOWABLE PRESSURE LOSS:
lowest design pressure and maximum allowable working
pressure (MAWP).
3. Set pressure is the pressure at which the relief valve starts to open,
SIZING RELIEF COND:
OVER PRESSURE USED:
(14.7 PSIA) at sea level.
5. Allowable overpressure is: 3% for fired steam boilers; 16% for
multiple valves; 21% for fire conditions; 10% for all other conditions.
6. Principal equipment where the PSV is located. Other equipment
to be protected is shown in the accompanying sketch.
7. The valve capacity at 10% overpressure which is the rated capacity
for API inlet and outlet hydraulic checks.
%
TOTAL API ORIFICE AREA REQ'D:
IN²
VALVE TYPE:
usually design pressure of equipment in PSIG.
4. Relief pressure is the set pressure + allowable overpressure +
5.55 PSI
RELIEF VALVE DATA
2. If more than one vessel is involved in a single relief system, use
DISCHARGE TO:
CONVENTIONAL
HEADER
BALANCED
ATM.
PILOT
PSV SIZE:
X
X
PROCESS
8 T 10
API ORIFICE AREA:
IN²
RATED FLOW
RATED FLOW IS CALCULATED AT 10% OVERPRESSURE
AND NO BACKPRESSURE FOR HYDRAULIC CHECKS
PSV SIZE:
API ORIFICE AREA:
T
IN²
RATED FLOW
Notice: This document has not been published and is the sole property of the Fluor Corporation, and is lent to the borrower for his confidential use only: and in consideration of the loan of this drawing, the borrower promises and agrees to return it upon request and agrees that it shall not be reproduced, copied, lent or otherwise disposed of directly or indirectly, nor used
for any purpose other than that for which it is specifically furnished.
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-25.xls]FIG D1.25-Data Sheet
FORM E-516i (REV 9/30)
Figure D1.26
No.
DATE
BY
REVISION
Sheet No.
Rev.
Date
15-Jun-2000
Manufacturer:
INSTRUMENT SPECIFICATIONS
PRESSURE SAFETY (RELIEF)
VALVES
P.O.
Unit:
Project:
Site:
Plant:
GENERAL
ALL ITEMS SHALL COMPLY WITH GENERAL INSTRUMENT SPECIFICATION SHEETS:
1 Tag Number
2 Service Description
3 Equipment No. Protected
4 P & ID / Section
Trim
5 Design Type
6 Bonnet
Nozzle
Lift
7 Conventional / Bellows / Pilot
BODY
8 Inlet:
Size - Rating / Facing
9 Outlet: Size - Rating / Facing
10 API Letter
Body / Bonnet Mat'l
11
TRIM
12 Nozzle
Seat / Disc
Guide / Rings
Pilot Trim
13 Spring
14 Bellows
15 Tube / Fittings / Sensor / Filter
MISC.
16 Cap: Threaded
18
19
ASME
Bolted
Plain
Packed
Vent With Bug Screen
Test Gag
Pilot Filter
B.Flo Prev
Test Con
ASME Code
UV-NB Stamps
Fire / P.F. / B.D. / C.W.F./ etc.
17 Lever:
20
21
22
23 Fluid
24 Required Capacity - Kg/Hr
25 Set Press-Kg/cm²G
26 Liquid SG
SIZING BASIS
27 Viscosity @ F.T., cP
28 Press-Kg/cm²G
29 Temp. °C
Nrm
%Overpressure
Gas MW
Latent Ht
Nrm
Max
Max
Relieving
30 Superimposed Back Pressure
31 Built Up Back Pressure
32 Compressibility, Z
Sp.Heat Ratio, k
33 Weight % Flash
34 With Rupture Disc
35 Failure Case
36 Discharges To
MFG DATA
37 Calculated * mm²
38 Spring Number*
39 Cold Diff Test Pressure
40 Leakage:
Selected
*
Adj. Range *
*
Other
Dim.
*
API 527
Capacity
42 Model Number
NOTES:
1. FURNISH A FIRMLY ATTACHED STAINLESS STEEL NAMEPLATE ENGRAVED WITH TAG NUMBER AND SERVICE.
TAG NUMBER SHALL BE STAMPED ALSO ON THE OUTLET FLANGE.
41 Certified: *
* - VENDOR INFORMATION AND / OR CONFIRMATION.
H:\SP\CENLIBRY\P000\225\Vol 48\[Figd1-26.xls]Sheet1
By
Chk'd
Appr'd
APPEJDIX
D-2
APPENDIX D
PAGE
1
TYPICAL CALCULATIONS
DATE
06-00
FLUOR DANIEL
FLARE SYSTEM
PROCESS MANUAL
APPENDIX D-2
TYPICAL CALCULATIONS
CALC
NUMBER
CHAPTER
CALCULATION TITLE
1
3
TYPICAL PSV SIZING, W/Back Pressure Correction
MANUAL
SECTION
NO.
3.1
2
3
SUBCRITICAL PSV SIZING
3.4
3
4
EMERGENCY DEPRESSURING
4.1.4
4
4
RELIEF LOAD SUMMARY
5
4
REBOILER PINCH
4.3.3.B
6
4
TOWER RELIEF LOADS
4.3.5
7
4
REACTOR LOOP RELIEF LOADS
4.4
8
4
SETTLING OUT PRESSURE
4.4.4
9
4
LOOP RELIEVING PRESSURE PROFILE
4.4.4.3
10
4
THERMAL RELIEF
4.5.2
11
4
RUPTURED TUBE RELIEF LOAD
4.7.1
12
4
DYNAMIC RELIEF ANALYSIS
4.7.3
13
4
FIRE RELIEF LOAD
4.8
14
4
FIRE DEPRESSURING LOAD
4.8.5
15
4
TANK VENT RATE
4.10
16
5
RISK BASED ASSESMENT
5.3.4
17
5
DYNAMIC SIMULATION
5.3.6
18
7
INLET LINE SIZING
7.2.1.5
19
7
OUTLET LINE SIZING
7.2.1.6
20
7
RELIEF HEADER SIZING
7.2.2
21
7
“IN-PLANT” RUN
7.3
22
8
BLOWDOWN DRUM SIZING
8.1
23
8
KO DRUM SIZING
8.2
24
8
SEAL DRUM SIZING
8.3
25
8
QUENCH DRUM SIZING
8.4
26
9
FLARE HEIGHT
9.4
27
9
EXPOSED SURFACE TEMPERATURES
9.4.6
28
9
SMOKELESS FLARING STEAM REQ’D
9.6
APXD2-R1.DOC