Flare System Process Design Manual

Flare System Process Design Manual PROCESS DESIGN MANUAL FLARE SYSTEM 670-225-9048 PLEASE READ MESSAGE BELOW BEFORE Y...

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Flare System Process Design Manual

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

PROCESS MANUAL 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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FLARE SYSTEM

PROCESS MANUAL 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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FLARE SYSTEM

PROCESS MANUAL 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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FLARE SYSTEM PROCESS MANUAL 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 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

Index-r1.doc

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


3.5.2

Applications

3.5.3


LIQUID TRIM RELIEF VALVES 3.6.1


3.6.2

Applications

3.6.3


3.7

SPECIAL FEATURES

3.8

PILOT OPERATED PRESSURE RELIEF VALVES

3.9

3.10

Index-r1.doc

3.8.1


3.8.2

Applications

3.8.3


3.8.4

Special Features

RUPTURE DISKS 3.9.1


3.9.2

Applications

3.9.3


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

Index-r1.doc

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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

Index-r1.doc

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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


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


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

Index-r1.doc

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


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


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)

Index-r1.doc

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9.0

FLARE

9.1

DESIGN DISCUSSION

9.2

9.3

9.4

9.5

9.6

Index-r1.doc

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


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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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PROCESS MANUAL INTRODUCTION

1.0

INTRODUCTION

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 Chap1-r1.doc

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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•

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 )] •


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 )]


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


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


Establish relief system design philosophy and utility failure basis.

E.) System Review


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


Secondary/Consultant Responsibility

Notes

A.) Basic Design


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


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


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



Secondary/consultant Responsibility

Notes

A.) Preliminary Relief Load Calculation


Control Systems Engineer

This step assists in identifying early needs for adequately sized PSVs and vessel nozzles

B.) Preliminary Valve Sizing



Process engineer performs sizing calculations.

C.) Preliminary Nozzle Size


Mechanical Engineer Piping Engineer

Unit Process Engineer coordinates with vessel group. Final nozzle size after relief device selection.

D.) Issue Preliminary PSV Data Sheets


Flare System Engineer Control System Engineer

Issue PSV data sheets to control system 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




Notes

A.) Flare Load Summary


Unit Process Engineers

Collect unit flare load summary with relief device calculation sheets. Check consistency.

B.) Preliminary Flare System Design


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


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


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.


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




Notes

7. Final Relief Load Computation and sizing


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



Issue final revision for Control Systems Engineer to prepare instrument data sheets and purchase specifications.

9. “As Purchased” Equipment Performance Review


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


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


11. Engineering Documentation

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Notes


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


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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ITEM


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


Compressed Air


Steam


Power


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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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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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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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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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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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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ITEM

2G. RELIEF VALVE CRITERIA


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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ITEM


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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ITEM


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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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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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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ITEM


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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ITEM


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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ITEM

5. HYDRAULICS 1. Flare header/subheader


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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PROCESS MANUAL DESIGN TEMPERATURE AND TEMPERATURE SELECTION

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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DESIGN TEMPERATURE SELECTION 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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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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RELIEVING DEVICES

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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3.2.4

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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3.2.6

RELIEVING DEVICES

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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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.


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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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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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Test Gags

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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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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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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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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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Filters

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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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Special Features

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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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o

RELIEVING DEVICES

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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2

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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


In which F2 is defined by: F2 = {[(k)/(k-1)][r

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][(1-r

(k-1)/k

)/(1-r)]}

0.5

3.5

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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

=


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 =


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

=


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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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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0.5

3.8 (Metric)

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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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PROCESS MANUAL DETERMINATION OF INDIVIDUAL RELIEF LOADS

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:

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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


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)


NT = fL/Di + K

4.21

0.5

Where: Gc

Chap4-r1.doc

=

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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Atmospheric Temperature Changes

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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

Chap4-r1.doc

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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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Atmospheric Pressure Changes

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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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Double wall refrigerated storage

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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.

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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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FLARE SYSTEM


FIGURE 4.1 ISOTHERMAL FLOW OF COMPRESSIBLE FLUIDS THROUGH PIPES AT HIGH PRESSURE DROPS

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OVERALL RELIEF SYSTEM LOAD EVALUATION

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OVERALL RELIEF SYSTEM LOAD EVALUATIONS

5.1

INTRODUCTION

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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 Chap5-r1.doc

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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. Chap5-r1.doc

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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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C)

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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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6.

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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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FIGURE 5.1 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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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.

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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.


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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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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. CHAP6-R1.DOC

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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.

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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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FIGURE 6.1 TYPICAL FLARE GAS RECOVERY SYSTEM

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FIGURE 6.2 FLARE GAS RECOVERY INLET PRESSURE CONTROL SYSTEM

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FIGURE 6.3 TYPICAL QUENCH DRUM

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FIGURE 6.4 TYPICAL SCRUBBER SYSTEM

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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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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


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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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:

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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)

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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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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:

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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.

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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: • •

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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)


Nonmod Pilot Nonmod Pilot Nonmod Pilot

VAP LIQ VAP/LIQ

50 (3) 50 (3) 50 (3)


Mod Pilot Mod Pilot Mod Pilot

VAP LIQ VAP/IQ

50 (3) 50 (3) 50 (3)


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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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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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PROCESS MANUAL

FIGURE 7.2 LABYRINTH TYPE SEAL

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PROCESS MANUAL

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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FLARE SYSTEM KNOCKOUT, BLOWDOWN, SEAL, QUENCH DRUMS AND PUMPS

PROCESS MANUAL

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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3

3

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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


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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o

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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:

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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:

•

•

•

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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.

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“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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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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PROCESS MANUAL FLARE

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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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


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)

CHAP9-R1.DOC

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

CHAP9-R1.DOC

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


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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)

CHAP9-R1.DOC

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)

CHAP9-R1.DOC

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


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)


9.30 (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 CHAP9-R1.DOC

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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


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 CHAP9-R1.DOC

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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).

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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

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APPENDIX A

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PROCESS MANUAL

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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PROCESS MANUAL

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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APPENDIX A

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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-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.

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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.

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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.

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APPENDIX B

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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.

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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.

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•

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.

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APPENDIX B

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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.

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•

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.

•


•


•


•


•


•

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.


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Feed Continues and Reflux Pump Continues (PPF-3): The Basis Factors (BF) are determined as follows: •


•


•


•


•


•



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.

•


•


These inputs result in a relief load of 0 lb/hr.

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APPENDIX B

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COOLING WATER FAILURE The Basis Factors (BF) are determined as follows: •


•


•


•


•

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


REFLUX FAILURE The Basis Factors (BF) are determined as follows: •


•


•


•


•

Loss of Reflux/Product pumps will cause the condensers to flood. CONDENSER BF = 0.00 and TRIM CONDENSER BF = 0.00.


REBOILER PINCH CALCULATION

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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.

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APPENDIX B

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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

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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

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FIGURE B1.1 PROCESS SKETCH

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BACKUP CALCULATIONS

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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.

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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

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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)

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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


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.


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


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.



4499920 A 26-Jun-2000 JLH DK

Appv'd:

DDC Page 4 of 12


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


1,6 1,3 4,6 5,6

1.00 0.00 1.00 0.00 0.00

TOTAL

144,908 583,395


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


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.



4499920 A 26-Jun-2000 JLH DK

Appv'd:

DDC Page 5 of 12


FEED REBOILER

BASIS FACTOR

1,4 2

1.00 1.00

TOTAL OUT:

728,303


CW FAILURE MW

M

86.6

°F 301.1

PSIA

BTU/LB

232.3

177.32


1,3 1,4 5 6

1.00 0.00 1.00 1.00 0.00

TOTAL

144,908 583,395


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


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.



4499920 A 26-Jun-2000 JLH DK

Appv'd:

DDC Page 6 of 12


FEED REBOILER

BASIS FACTOR

1,3 2

1.00 1.00

TOTAL OUT:

728,303


REFLUX FAILURE MW

M

86.6

°F 301.1

PSIA

BTU/LB

232.3

177.32


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


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.



4499920 A 26-Jun-2000 JLH DK

Appv'd:

DDC Page 7 of 12


FEED REBOILER


1.00 1.00

TOTAL OUT:

728,303

M

86.6

301.1

232.3

177.32


1,6 1,3 4,6 5,6

1.00 0.00 1.00 0.63 0.50

TOTAL

144,908 583,395


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


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.



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


LB/HR

0.00 0.00

0

TOTAL OUT:

M

86.6

301.1

232.3

BTU/HR

177.32

0 0


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


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


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.



4499920 A 26-Jun-2000 JLH DK

Appv'd:

DDC Page 9 of 12

BTU/LB

BTU/HR


BASIS FACTOR

FEED REBOILER

0

TOTAL OUT:


LB/HR

MW

M

86.6

°F 301.1

PSIA 232.3

177.32


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


0.0

MESSAGE: BASIS:

NOTES:

0 0



4499920 A 26-Jun-2000 JLH DK

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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:

DK

Appv'd:

DDC 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

Revision :

A

Date :

26-Jun-2000

By:

JLH

Chk:

DK

Appv'd:

DDC Page 12 of 12

ACCUMULATOR SURGE TIME PLANT: RFCC-IRVING OIL EQ NO: C-42517


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


Feed Stops and Reflux Pump Stops

Feed Continues and Reflux Pump Stops

REFLUX FAILURE

CW FAILURE


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


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


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


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


4499920 A 15-Jun-2000 JLH DK

Appv'd:

DDC 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


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


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


MESSAGE: BASIS:

NOTES:


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

Contract :

04411700

Revision :

A

Date :

15-Jun-2000

By:

Figure B4.1

Your Name Here

Chk:

Appv'd:

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


1.2735 BTU/SCF

2,575

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 = 2.51/(C_L^0.625)

4.63

4.63

X_L

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 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


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.



4499920 A 23-Jun-2000 JLH DK

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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


0.00 1.00

TOTAL OUT:

LB/HR 0

M

86.6

301.1

232.3

177.32


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


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


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.



4499920 A 23-Jun-2000 JLH DK

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DDC Page 4 of 12

Figure D1.4


FEED REBOILER


1.00 1.00

TOTAL OUT:

728,303

M

86.6

301.1

232.3

BTU/LB 177.32


1,6 1,3 4,6 5,6

1.00 0.00 1.00 0.00 0.00

TOTAL

144,908 583,395


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


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.



4499920 A 23-Jun-2000 JLH DK

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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


CW FAILURE MW

M

86.6

°F 301.1

PSIA

BTU/LB

232.3

177.32


1,3 1,4 5 6

1.00 0.00 1.00 1.00 0.00

TOTAL

144,908 583,395


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


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.



4499920 A 23-Jun-2000 JLH DK

Appv'd:

DDC Page 6 of 12

Figure D1.6


FEED REBOILER

BASIS FACTOR

1,3 2

1.00 1.00

TOTAL OUT:

728,303


REFLUX FAILURE MW

M

86.6

°F 301.1

PSIA

BTU/LB

232.3

177.32


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


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.



4499920 A 23-Jun-2000 JLH DK

Appv'd:

DDC Page 7 of 12

Figure D1.7


FEED REBOILER


1.00 1.00

TOTAL OUT:

728,303

M

86.6

301.1

232.3

177.32


1,6 1,3 4,6 5,6

1.00 0.00 1.00 0.63 0.50

TOTAL

144,908 583,395


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


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.



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


LB/HR

0.00 0.00

0

TOTAL OUT:

M

86.6

301.1

232.3

BTU/HR

177.32

0 0


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


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


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.



4499920 A 23-Jun-2000 JLH DK

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DDC Page 9 of 12

BTU/LB

BTU/HR

Figure D1.9


BASIS FACTOR

FEED REBOILER

0

TOTAL OUT:


LB/HR

MW

M

86.6

°F 301.1

PSIA 232.3

177.32


0 0

TOTAL

0 V L L

50.5

178.4

229.3

207.09

105.3

435.7

234.3

222.11

0


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


0.0

MESSAGE: BASIS:

NOTES:

0 0



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


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


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



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 -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


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


1.2735 BTU/SCF

2,575

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 = 2.51/(C_L^0.625)

5.15

5.15

X_L

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 0.45, Z/L = 0.09. If Ua/Ug 0.05 and Ua/Ug 0.45, X/L = 0.97. If Ua/Ug 0.05 and Ua/Ug

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