119,99 €
Mehr erfahren.
- Herausgeber: John Wiley & Sons
- Kategorie: Fachliteratur
- Sprache: Englisch
A practical and engaging guide to running process controls in petrochemical plants and refineries
Process control is an area of study dealing with controlling variables that emerge in process plants, such as chemical plants, wastewater purification plants, or refineries. Existing guides to process control are numerous, but they tend to be associated with control engineering, which is more mathematical and theoretical. There is an urgent need for a more straightforward and concrete guide for practical use in petrochemical plants and refineries.
Troubleshooting Process Plant Control meets this need with a work dedicated to real-life solutions and problem solving. Rooted in real-world examples and the career experience of the author, it largely avoids complex mathematics in favor of practical, well-established process engineering principles. Now fully updated to reflect the latest best practices and developments in the field, it is indispensable for process controllers in active plants of all kinds.
Readers of the third edition will also find:
- New chapters on alarm disabling, spectrometer use, and reducing CO2 emissions
- Additional novel examples throughout
- Guidelines for using spectrometers to directly control reflux rates and steam flow to reboilers
Troubleshooting Process Plant Control is ideal for practicing engineers and other technical professionals working in process facilities, as well as advanced students taking professional training courses in these fields.
Sie lesen das E-Book in den Legimi-Apps auf:
Seitenzahl: 532
Veröffentlichungsjahr: 2024
Ähnliche
Table of Contents
Cover
Table of Contents
Series Page
Title Page
Copyright Page
Dedication to Third Edition
About the Author
Preface to the Third Edition
Introduction
Auto Versus Manual
Gain, Reset, and Advanced Computer Control
1 Learning from Experience
Learning From Plant Operators
Learning From Field Observations
Learning From Mistakes
Learning From Theory
Learning From Relationships
Normal Purpose of Control Loops
Manual Versus Auto
Process Control Nomenclature
2 Process Control Parameter Measurement
How are Liquid Levels Measured?
How are Temperatures Measured?
External Fouling of Thermowell
Furnace Temperatures
Measuring Pressures
Measuring Differential Pressures
How are Flows Measured?
3 Dependent and Independent Variables
Changing the Degrees of Freedom
Variables in Distillation
Increasing the Degrees of Freedom
Complex Distillation Towers
Amoco‐Texas City—1974
Variables in Gas Compression
4 Binary Distillation of Pure Components
Proper Control of Water Strippers
5 Distillation Tower Pressure Control
Cooling Water Throttling
Pressure Control with Noncondensable Vapors
Flooded Condenser Pressure Control
Hot Vapor Bypass Pressure CONTROL
Combining Noncondensable and Total Condensation
Split‐Range Control for Noncondensables
Reboiler Controls Tower Pressure
Tower Top Pressure Sensing Point
6 Control of Aqueous Phase (Waste Water) Strippers
Controlling Amine Regenerator and Sour Water Stripper Reboiler Steam Flows
Controlling Reflux Rate
Use of Oil Skimming Site Glass
Controlling Stripping Steam Rate on Towers with No Reflux
Required Stripper Efficiency
Control of Stripper Pressure
Safety Note
7 Pressure Control in Multicomponent Systems
Getting a Sample in a BOTTLE
Heat Does Not Always Make Pressure
Effect on LPG Quality
Test Question
Refrigerant Composition
Adjusting Refrigerant Composition
Explaining Multicomponent Refrigeration to the Panel Board Operators
8 Optimizing Fractionation Efficiency by Temperature Profile
Advanced Computer Control
Superadvanced Computer Control
Full Flood
The Case for Feed PREHEAT
9 Analyzer Process Control
Controlling Diesel Draw‐Off Rate
Direct Analyzer Control
De‐Ethanizer Reboiler Control
Control of Asphalt Viscosity
Example of Direct Analyzer Process Control
Direct Analyzer Control
Spectrometer Control
10 Fired Heater Combustion Air Control
Control of Air with O
2
Analyzer
Setting a Target
The Point of Absolute Combustion as a Variable
Automatic Temperature Control
Combustible Analyzer Control
Correct Control of Combustion Air
Fired Heater Safety Note
11 Using Existing Controls to Promote Energy Efficiency
Compression Energy Savings
Elimination of Orifice Plate Pressure Drop
Undersized Control Valves Waste Energy
Opening Control Valve Bypasses
Hand‐Jacks
Oversizing Control Valves
12 Sizing Process Control Valves
Sizing Control Valves
Undersizing Control Valves
Why Ignore Changes in Elevation and Operating Pressure?
Energy Losses In Control Valves
Checking Control Valve Capacity
Increasing Control Valve Size
Effect of Oversizing Control Valves
Unresponsive Control Valves
13 Control Valve Position on Instrument Air Failure
Instrument Air Pressure Control Box
Control Valve on Discharge oF Pump
Control Valve on the Discharge of a Compressor
Pressure Control of Vessels
Fuel Gas to Heaters
Avoid Loss of Air Pressure
Control Valve Chatter & Leakage
14 Override and Split‐Range Process Control
Cascade Control
Override Control
Split‐Range Control
How Does Split‐Range Control Work?
Safety Tips
Enhanced Control Valve Safety
15 Vacuum System Pressure Control
Factors Affecting Loss of Sonic Boost
Specialty Vacuum Ejectors
Throttling on Motive Steam
Positive Feedback Loop
Spillback Pressure Control
Varying Tower Top Pressure
Summary of Control of Vacuum Systems
Spillback
Throttling Motive Steam to First‐Stage Ejector
Throttling on Inlet to First‐Stage Jet
16 Reciprocating Compressors
High Discharge Temperature Trip
Measuring Process Flows
Flow Control
17 Centrifugal Compressor Surge versus Motor Over‐Amping
Understanding Blower Controls
The Second Law of Thermodynamics
Effect of Wet Gas Molecular Weight
Summary
18 Controlling Centrifugal Pumps
Pump Suction Pressure Versus Level Control
Determining Suction Set Point Pressure
Turbine‐Driven Pumps
Safe Minimum Flow Control
Undersized Control Valve Reduces Pump Capacity
19 Steam Turbine Control
Steam Turbine Theory
Use of The Hand Valves
Turbine Over‐Speed Trip
Turbine Exhaust—Surface Condenser Level Control
Level Control Malfunction
Controlling Moisture Content of Turbine Exhaust Steam
20 Steam and Condensate Control
Condensate Level Control
Condensate Drum
Condensate Flow Problems
Boiler Level Control Causes CARRYOVER
Steam Flow to a Reboiler
21 Control of Process Reactions
Control of Hydrotreater Hydrogen Recycle Flow
Backup System for Control oF Air to the Mercaptan Sweetening Plant
Unstable Recycle Gas Flow on a Vacuum Gas Oil Hydrotreater
Control of Fluid Cracking Catalyst Activity
22 Function of the Process Control Engineer
Process Control Engineer’s Safety Responsibility
Later that Evening
Scope of Control Engineer’s Job
Bogging Down a Process Heater
23 Steam Quality and Moisture Content
Flowing Steam
Why Boilers Carry Over
Level Control in a Kettle Waste Heat Boiler
Overflow Baffle in Kettle Waste Heat Boiler
Level Control in Deaerators
Controlling Water Content of Stripping Steam
24 Level, Pressure, Flow, and Temperature Indication Methods
Effects of Temperature on Level
Plugged Taps
High Liquid Level
Effect of Aerated Liquid on Level Indications
Split Liquid Levels
Radiation Level Detection
Vacuum Pressure Measurements
Flow Indication
Non‐Orifice‐Type Flow Measurement Methods
Checking Flows in the Field
Correcting Flowmeter Reading Off‐Zero
Temperature Measurement
References
25 Alarm and Trip Design for Safe Plant Operations
The Concept of Redundancy
Testing the New Pressure Alarm
Use of Conductivity Probe
High‐Temperature Trips
Flows
Optical Sensors
Flushing of Connections
How Not to Test a TRIP
26 Inverted Response of Process Parameters
Effect of the Top Tray Flooding
Common Examples of Inverted Temperature Response
Centrifugal Pump Power Requirement
Variable‐Speed Turbine Drive
Hot Vapor Bypass Pressure Control
27 Nonlinear Process Responses
Effect of Nozzle Exit Loss on Flow Indication
Nonlinear Liquid Level Indications
Centrifugal Pump Discharge Pressure
28 Control Malfunction Stories
High‐Pressure Alarm
Excess Moisture in Turbine Steam
Thermocouple Length
Coked‐Over Thermowell
29 Level Indication Problems in Vessels
Improper Level Tap Location
Elevation of Stripping Steam Nozzle
Moisture in Top Level Tap
Radiation Level Detection
Optimization of Liquid Level Tap Locations
The Purpose of Vessel Level Control
Observing Centrifugal Pump Cavitation
30 Calibration Specific Gravity for Level Control
Selecting Calibration S.G.
Becoming “Tapped‐Out”
Identifying When a Vessel Becomes Tapped‐Out
31 Flow Orifice Plate Cavitation
Field Investigation
Problem Resolved
A Lesson Learned
32 Factors Causing Incorrect Process Parameter Measurements
Temperature
Pressure Indication
Location of a Vessel High‐Pressure Alarm Indicator
Level Indication
Calibration Specific Gravity
Moisture Contamination of Hot Oil Level Indication
Flows
Composition
Lab Analysis Problems
Analyzer Closed‐Loop Control
Reference
33 Optimizing Controls to Reduce Emissions from Sulfur Recovery Plants and Sour Water Strippers
Naphtha in Sour Water
Origins of Naphtha in Sour Water Stripper Feed
Sulfur Plant Air Consumption
Sulfur Plant Air Deficiency
CO
2
Emissions
Stripping Steam Usage
Monitoring Naphtha in Sour Water Stripper Bottoms
On‐Stream Analysis of Stripper Bottoms
Reference
34 Positive Pressure in Fired Heaters—Effect of Wind
Why is a Positive Pressure Never Permitted
Development of a Positive Pressure
Effect of Wind
Effect of Wind Gusts 60 Feet above the Bottom of the Heater
Effect of Wind on Combustion Air Flow
Basis for Above Observations
35 A Lesson From Bhopal—Hazards of Ignoring Alarms
Phosgene
The Problem in Bhopal
Pump Mechanical Seal Leakage
Alarms Disabled
Crystal Formation
Phosgene Vented to the Atmosphere
About My Seminars
What’s NEW
Synergism
Who Caused The Problem?
Conservation Ideas In The SEMINAR
Further Readings on Troubleshooting Process Controls
The Norm Lieberman Video Library of Troubleshooting Process Operations
Process Control Nomenclature Used in Petroleum Refineries and Petrochemical Plants
Index
End User License Agreement
List of Illustrations
Chapter 1
Figure 1‐1 Adjusting wash oil based on gas oil color
Figure 1‐2 Opening spillback to keep the FRC valve in its linear operating r...
Figure 1‐3 Unintentional flaring caused by malfunction of the LPG makeup con...
Figure 1‐4 Tuning a level control valve depends on what is downstream
Figure 1‐5 Too much steam flow causes a loss in vacuum
Chapter 2
Figure 2‐1 Measuring levels by sensing liquid head pressure
Figure 2‐2 The thermocouple wire should be fully inserted in the thermowell ...
Figure 2‐3 Never connect an alarm and control to the same sensing point
Figure 2‐4 Flows are measured by inducing a delta P through an orifice
Chapter 3
Figure 3‐1 Simple distillation tower
Figure 3‐2 Simple refrigeration circuit
Chapter 4
Figure 4‐1 Controlling steam flow to a stripper reboiler
Figure 4‐2 Rich amine flow controls steam to the reboiler to hold a fixed ra...
Figure 4‐3 Reboiler steam controlled by reflux, reset by feed rate
Chapter 5
Figure 5‐1 Pressure fluctuations due to the control valve operating in its n...
Figure 5‐2 Control of cooling water flow for pressure control is not recomme...
Figure 5‐3 Flooded condenser pressure control incorrectly applied
Figure 5‐4 Flooded condenser pressure control is correct design practice
Figure 5‐5 Hot vapor bypass pressure control is wrong design practice
Figure 5‐6 Flooded condenser pressure control with provision for noncondensa...
Figure 5‐7 Split‐range pressure control using makeup gas when vent valve shu...
Figure 5‐8 Heat input directly controlling tower pressure is a correct desig...
Chapter 6
Figure 6‐1 Control reboiler steam to maintain a desired flux rate and not th...
Figure 6‐2 Operator raises level and observes if material flowing through th...
Chapter 7
Figure 7‐1 Gasoline stabilizer with condenser capacity limiting pressure con...
Figure 7‐2 No. 2 alkylation unit sets record, March, 1975 at 23,851 B/D. Aut...
Figure 7‐3 Maximizing refrigeration capacity by override control
Chapter 8
Figure 8‐1 Optimizing feed preheat
Figure 8‐2 Override control maximizes tower delta T
Chapter 9
Figure 9‐1 Maximizing diesel production with an online analyzer
Figure 9‐2 Heat input directly controlled by product spec
Figure 9‐3 Control stripping steam for asphalt viscosity specifications
Figure 9‐4 Sulfur recovery and tail gas unit
Figure 9‐5 Sample sketch of the tec5USA Raman Immersion Probe installation t...
Chapter 10
Figure 10‐1 Optimizing air flow at constant fuel
Figure 10‐2 The point of absolute combustion defines the optimum combustion ...
Figure 10‐3 Analyzer sets air flow. This is a bad design
Figure 10‐4 Correct strategy to control combustion air
Chapter 11
Figure 11‐1 Eliminating DP of orifice place saves energy
Chapter 12
Figure 12‐1 A poorly designed control valve installation
Figure 12‐2 Valve position shown 75% open
Figure 12‐3 The 4″ × 1″ piping reducer to accommodate small control va...
Chapter 13
Figure 13‐1 A control valve that will fail closed on loss on of instrument a...
Figure 13‐2 There are three air pressures shown on the box adjacent to the c...
Figure 13‐3 Level control valve fails safely in a closed position
Figure 13‐4 Pressure control valve failure position depends on feed control ...
Figure 13‐5 Example of air failure valve positions for heater
Figure 13‐6 Example of air failure valve positions for a distillation tower...
Chapter 14
Figure 14‐1 Override pressure control on a boiler feed water deaerator
Figure 14‐2 Split‐range pressure control of a distillation column
Chapter 15
Figure 15‐1 Converging–diverging vacuum steam ejector
Figure 15‐2 Throttling on motive steam supply sometimes works well in vacuum...
Figure 15‐3 Spillback pressure control. Too many jets in service
Figure 15‐4 Three options to control vacuum tower pressure
Chapter 16
Figure 16‐1 High differential pressure trips the compressor on high discharg...
Figure 16‐2 Flow control with spillback is an energy—wasting design
Figure 16‐3 Piston head shown at end of its travel
Chapter 17
Figure 17‐1 Filter plugging reduces motor amps on the sulfur plant air blowe...
Figure 17‐2 Fixed‐speed centrifugal compressor operating curve. Polytropic h...
Figure 17‐3 Suction throttling pressure control for a motor‐driven centrifug...
Figure 17‐4 Spillback suction pressure control. Constant‐speed compressor...
Chapter 18
Figure 18‐1 Suction pressure control with fixed—speed pump. Level control no...
Figure 18‐2 Suction pressure control with variable—speed pump. Level control...
Figure 18‐3 Protecting the pump from too low a flow with minimum‐flow “Yarwa...
Chapter 19
Figure 19‐1 Steam turbine component functions
Figure 19‐2 Turbine‐driven process pump
Figure 19‐3 Direct speed control by process parameter
Figure 19‐4 An air leak in the gauge glass caused condensate backup in the s...
Chapter 20
Figure 20‐1 Low‐pressure condensate flows erratically into a collection head...
Figure 20‐2 Condensate level controls flow of steam into the exchanger but a...
Figure 20‐3 Properly instrumented condensate level control for stability. No...
Figure 20‐4 Flashing condensate causes backup in the steam heater
Figure 20‐5 Boiler level control causes carryover
Chapter 21
Figure 21‐1 Delayed coking is a zero‐order endothermic reaction
Figure 21‐2 Hydrodesulfurization unit with recycled H
2
Figure 21‐3 Mercaptan sweetening plant
Figure 21‐4 Unstable recycle gas flow on a gas oil hydrodesulfurization unit
Figure 21‐5 Centrifugal compressor performance curve
Chapter 22
Figure 22‐1 Backup on low lube oil pressure should be motor‐driven pump...
Figure 22‐2 Three‐position switch governs control of the backup
Figure 22‐3 Forgotten control limit cost Texaco $5,000,000
Chapter 23
Figure 23‐1 Hydrogen plant waste heat boiler in Aruba
Figure 23‐2 Steam enthalpy versus entropy
Figure 23‐3 TI and PI used to determine level set point
Figure 23‐4 Level control in a kettle waste heat steam generator
Figure 23‐5 Baffles in kettle waste heat boilers are design errors and serve...
Figure 23‐6 Level control problem in deaerator. Cascade control of pressure ...
Figure 23‐7 Override pressure control on a boiler feed water deaerator
Figure 23‐8 Stripping steam inlet to low‐pressure large‐diameter towers to p...
Chapter 24
Figure 24‐1 A gauge glass functions as a manometer
Figure 24‐2 The circulation of liquid in a gauge glass
Figure 24‐3 Operation of a level‐troll
Figure 24‐4 Split liquid level indication caused by foam
Figure 24‐5 Foam creates a nonlinear response in level indication
Figure 24‐6 A mercury absolute‐pressure manometer
Figure 24‐7 Orifice flowmeter
Figure 24‐8 Ultrasonic transit time flow measurement technology.
Chapter 25
Figure 25‐1 High‐pressure alarm senses lower pressure from the same connecti...
Figure 25‐2 Improper level instrumentation
Figure 25‐3 Corrected level instrumentation
Figure 25‐4 Level alarm or trip configuration
Figure 25‐5 Flow trip or alarm configuration
Chapter 26
Figure 26‐1 Opening the condensate drain valve too much leads to a large los...
Figure 26‐2 My propane–butane splitter in Texas City
Figure 26‐3 Response of amp load on the motor driver by throttling on discha...
Figure 26‐4 Closing the hot vapor bypass control valve may increase instead ...
Chapter 27
Figure 27‐1 Adjusting wash water to prevent HCl corrosion in the exchanger
Figure 27‐2 Finding the forced condensation dew point temperature
Figure 27‐3 Flow indication rises as level falls due to nozzle exit loss cav...
Figure 27‐4 Effect of flashing liquid entering an orifice plate
Figure 27‐5 Foam and liquid in the vessel, but only stagnant liquid in level...
Figure 27‐6 Nonlinear level response of liquid level to foam
Chapter 28
Figure 28‐1 False low level indication
Figure 28‐2 Pressure control and high‐pressure alarm on the same ¾‐inch conn...
Figure 28‐3 Opening a hand valve will reduce moisture content in exhaust ste...
Chapter 29
Figure 29‐1 Never locate top level tap above seal pan
Figure 29‐2 Never locate steam inlet on same side and below stripping steam ...
Chapter 30
Figure 30‐1 Level controller actually measures a pressure difference between...
Chapter 31
Figure 31‐1 Mislocated orifice plate causes butane‐splitter flaring
Chapter 33
Figure 33‐1 Sour water stripper
Figure 33‐2 tec5USA online composition analyzer
Chapter 34
Figure 34‐1 Natural draft fired heater
Chapter 35
Figure 35‐1 Water wash by‐passed pump discharge check‐valve
Guide
Cover Page
Table of Contents
Series Page
Title Page
Copyright Page
Dedication to Third Edition
About the Author
Preface to the Third Edition
Introduction
Begin Reading
About My Seminars
Further Readings on Troubleshooting Process Controls
The Norm Lieberman Video Library of Troubleshooting Process Operations
Process Control Nomenclature Used in Petroleum Refineries and Petrochemical Plants
Index
WILEY END USER LICENSE AGREEMENT
Pages
ii
iii
iv
v
ix
xi
xii
xiii
xiv
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
22
23
25
26
27
28
29
30
31
32
33
34
35
36
37
38
39
41
42
43
44
45
46
47
48
49
50
51
52
53
54
55
56
57
58
59
60
61
62
63
64
65
66
67
68
69
70
71
72
73
74
75
76
77
78
79
80
81
82
83
84
85
86
87
88
89
90
91
92
93
94
95
97
98
99
100
101
102
103
105
106
107
108
109
110
111
112
113
114
115
116
117
118
119
120
121
122
123
124
125
126
127
128
129
130
131
132
133
134
135
136
137
138
139
141
142
143
144
145
146
147
148
149
150
151
152
153
154
155
156
157
158
159
160
161
162
163
164
165
166
167
168
169
170
171
172
173
174
175
177
178
179
180
181
182
183
184
185
186
187
188
189
190
191
192
193
194
195
196
197
198
199
200
201
203
204
205
206
207
208
209
210
211
213
214
215
216
217
218
219
220
221
222
223
224
225
226
227
228
229
230
231
232
233
234
235
236
237
238
239
240
241
242
243
245
246
247
248
249
250
251
252
253
254
255
256
257
258
259
260
261
262
263
264
265
266
267
268
269
271
272
273
274
275
276
277
278
279
280
281
282
283
284
285
286
287
288
289
290
291
292
293
294
295
296
297
298
299
300
301
302
303
305
306
307
308
309
310
311
312
313
314
315
319
321
322
323
324
325
327
328
329
330
331
332
333
334
335
336
337
338
339
340
341
342
343
344
345
346
347
348
349
350
351
352
353
354
355
Troubleshooting Process Plant Control
Other Books by Norman P. Lieberman
Troubleshooting Refinery Operations
—PennWell Publications
Troubleshooting Process Operations 5th Edition
—PennWell Publications
A Working Guide to Process Equipment
(with E. T. Lieberman)—5th Edition—McGraw Hill Publications
Troubleshooting Natural Gas Processing
(order by e‐mail at
norm@lieberman‐eng.com
)
Process Design for Reliable Operations 3rd Edition
(order by e‐mail at
norm@lieberman‐eng.com
)
Troubleshooting Vacuum Systems
—John Wiley & Sons Publications.
Process Engineering for a Small Planet—
John Wiley & Sons Publications.
Process Equipment Malfunctions
—McGraw Hill Publications
Process Engineering:
Facts, Fiction, and Fables—John Wiley
Understanding Process Equipment for Operators and Engineers
—Elsevier
Process Operations:
Lessons Learned in a Nontechnical Language—PennWell Publications
My Race with Death
(order by e‐mail at
norm@lieberman‐eng.com
)
Copies of the first three texts are best ordered from the publishers, but may be ordered through us. E‐mail (norm@lieberman‐eng.com). Troubleshooting Refinery Operations (1980) has been incorporated into Troubleshooting Process Operations.
Troubleshooting Process Plant Control
A Practical Guide to Avoiding and Correcting Mistakes
Third Edition
Norman P. Lieberman
Chemical EngineerProcess Improvement EngineeringMetairie, Louisiana, USA
Copyright © 2024 by John Wiley & Sons, Inc. All rights reserved, including rights for text and data mining and training of artificial technologies or similar technologies.
Published by John Wiley & Sons, Inc., Hoboken, New Jersey.Published simultaneously in Canada.
No part of this publication may be reproduced, stored in a retrieval system, or transmitted in any form or by any means, electronic, mechanical, photocopying, recording, scanning, or otherwise, except as permitted under Section 107 or 108 of the 1976 United States Copyright Act, without either the prior written permission of the Publisher, or authorization through payment of the appropriate per‐copy fee to the Copyright Clearance Center, Inc., 222 Rosewood Drive, Danvers, MA 01923, (978) 750–8400, fax (978) 750–4470, or on the web at www.copyright.com. Requests to the Publisher for permission should be addressed to the Permissions Department, John Wiley & Sons, Inc., 111 River Street, Hoboken, NJ 07030, (201) 748–6011, fax (201) 748–6008, or online at http://www.wiley.com/go/permission.
Trademarks: Wiley and the Wiley logo are trademarks or registered trademarks of John Wiley & Sons, Inc. and/or its affiliates in the United States and other countries and may not be used without written permission. All other trademarks are the property of their respective owners. John Wiley & Sons, Inc. is not associated with any product or vendor mentioned in this book.
Limit of Liability/Disclaimer of Warranty: While the publisher and author have used their best efforts in preparing this book, they make no representations or warranties with respect to the accuracy or completeness of the contents of this book and specifically disclaim any implied warranties of merchantability or fitness for a particular purpose. No warranty may be created or extended by sales representatives or written sales materials. The advice and strategies contained herein may not be suitable for your situation. You should consult with a professional where appropriate. Further, readers should be aware that websites listed in this work may have changed or disappeared between when this work was written and when it is read. Neither the publisher nor authors shall be liable for any loss of profit or any other commercial damages, including but not limited to special, incidental, consequential, or other damages.
For general information on our other products and services or for technical support, please contact our Customer Care Department within the United States at (800) 762–2974, outside the United States at (317) 572–3993 or fax (317) 572–4002.
Wiley also publishes its books in a variety of electronic formats. Some content that appears in print may not be available in electronic formats. For more information about Wiley products, visit our web site at www.wiley.com.
Library of Congress Cataloging‐in‐Publication Data
Names: Lieberman, Norman P., author. | John Wiley & Sons, publisher.Title: Troubleshooting process plant control : a practical guide to avoiding and correcting mistakes / Norman P. Lieberman.Description: Third edition. | Hoboken, New Jersey : Wiley, [2024] | Includes index.Identifiers: LCCN 2024016445 (print) | LCCN 2024016446 (ebook) | ISBN 9781394262939 (hardback) | ISBN 9781394262953 (adobe pdf) | ISBN 9781394262946 (epub)Subjects: LCSH: Petroleum refineries–Maintenance and repair. | Chemical process control.Classification: LCC TP690.3 .L534 2024 (print) | LCC TP690.3 (ebook) | DDC 660/.2815–dc23/eng/20240513LC record available at https://lccn.loc.gov/2024016445LC ebook record available at https://lccn.loc.gov/2024016446
Cover Design: WileyCover Image: © Chanin Nont/Getty Images
Dedication to Third Edition
Time goes on. And life goes on. The seasons progress from fall, to winter, into spring, same with me. I was young, then progressed to middle age. Became old. And then I peaked at age 82, and became younger again!
How did this happen?
My wife and inspiration, Liz, explained it to me, “Norm, you had a chance to retire ten years ago. Now it’s too late.”
I guess then I’ll have to go on to the end, Liz and I together. Old process engineers never die; they just fade away.
March, 2024
About the Author
Norman P. Lieberman has been troubleshooting refinery and chemical plant process equipment since 1964. He began work at American Oil as a process engineer at their Indiana Refinery. Lieberman designed and operated sulfur plants, alkylation units, LPG fractionators, cokers, distillation towers, and vacuum systems for Amoco until 1980. He was next employed at the Good Hope Refinery in New Orleans where he worked on their polymerization unit, MTBE plant, crude unit, naphtha reformer, and hydrodesulfurization facilities.
In 1985, Lieberman worked for Good Hope, troubleshooting gas field compression, dehydration, and treating problems for their Loredo, TX natural gas production facilities. In 1988, he became a consultant for petrochemical and refinery process equipment problems.
Approximately 24,000 engineers and operators have attended his 1030 in‐house troubleshooting seminars and webinars. Lieberman has a degree in Chemical Engineering, 1964, from Cooper Union. He lives in New Orleans with his wife, Liz, also a chemical engineer.
Norm runs every day for three miles and plans to do so until the end.
Preface to the Third Edition
I have been practicing process engineering for 59 years. Mainly, as a refinery field troubleshooter for distillation operations, vacuum systems, fired heaters, waste water strippers, compressors, and pumps. The majority of the malfunctions I discover are not due to faulty equipment design, mechanical failure, or operator error. The big problem is with process control and measurement of process variables.
I always explain during my process equipment troubleshooting seminars, which I’ve instructed since 1983, that the process control engineer is the most important person in the plant. I was sure of that in 1983 and am equally sure as I write these words in 2024.
The problem that the refining, petrochemical, and chemical fertilizer industry has is that the university course of study for process control engineers is worse than bad. It’s irrelevant! In 1979 at Northwestern, and in 1983 at LSU, I found this out personally, having been ejected from both institutions after 1 day as an instructor. My conception of the training required to be an effective process control engineer, being at odds with that of both universities.
Process control has little to do with math, or computers, or Laplace transforms. It’s about understanding the following:
How instruments work.
How variables of temperature, pressure, level, flows, and composition are measured in the field.
How controls interact with process equipment.
How unit operators interact with controls.
The tendency of instrumentation to be trapped in a “positive feedback loop.”
How process equipment itself works from a chemical engineering perspective.
How to make field measurements.
I wasn’t particularly knowledgeable about process control until 1974, even though I had been employed by Amoco Oil for 10 years. But in 1974, I worked as an operator for 4 months, during a strike in Texas City. Then again in 1980, there was an even longer strike, part of which I worked as the panel operator on a sulfur recovery and amine unit. Afterward, I was fairly competent to tackle a variety of process control issues.
Based on my subsequent 44 years of experiences, I have developed the following advice for young process control engineers:
“The price we pay for success is the willingness to risk failure.” Michael Jordan, Chicago Bulls.
You can email me with questions at norm@lieberman‐eng.com
Introduction
Troubleshooting process plant control first requires an understanding of a wide variety of malfunctions that may develop in measuring variables such as:
Levels
Temperatures
Pressures
Flows
Compositions
Certainly, if you cannot measure a level or pressure, you can’t expect to control it.
Second, the console operator or process engineer must understand how the control valves and the signals to open and close the valves actually work. For instance, did you know that the valve position shown in the control center does not at all represent the actual valve position? It represents what the control valve position is supposed to be.
To troubleshoot process control problems, the operator, or engineer, has to understand the relationship of the controls to the individual process. This means, you will have to get to know the unit and how it works. This is the most difficult part of the job of understanding process plant control.
Many apparent control problems are, in reality, process problems. But, on the other hand, after 59 years of troubleshooting refinery process problems, I am quite sure that the most common sort of malfunctions I have encountered are related to the inability to measure a level, temperature, or flow correctly—and also to have a control valve respond in the manner needed to achieve the desired operational change.
AUTO VERSUS MANUAL
When you see that the console operator is running a control loop on “manual” rather than in “auto,” that is an indication that something is wrong with the field measurement of the variable, or with the control logic. Is it a metering problem, a sensor that is fouled, or a variable that is over‐ranged? Perhaps, the variable is caught up in a “positive feedback loop”? Control loops are supposed to run in auto, and you should not accept loops that run in manual, as representing an acceptable mode of operation. Sooner or later, such broken (i.e., manual) control loops will slip out of an acceptable operating range.
GAIN, RESET, AND ADVANCED COMPUTER CONTROL
This text does not deal with the time aspects of optimizing the relationship between variables. Typically, the console operator is far less concerned about how fast operating parameter returns to its set point, than if a particular variable is moving in the right direction, so that he can safely bring his products back on spec.
Advanced computer control is largely irrelevant to my work in field troubleshooting refinery and petrochemical plant process control problems. I cannot conceive as to why process control engineering is so often taught in universities as if it is a form of higher mathematics. Even more detrimental to unit operations is the perception by plant management and staff engineers that advanced computer controls are actually being utilized on the operating units, when in reality the console operators are struggling to run critical control loops on auto, without getting caught up in a dangerous positive feedback loop. I never understood anything about Laplace transforms in school, and I am certainly too old to start learning now.
The real problems with hydrocarbon processing that I first encountered in 1965 are 90% unchanged in my work as a chemical engineer in 2024.
1Learning from Experience
An old Jewish philosopher once said, “Ask me any question, and if I know the answer, I will answer it. And, if I don’t know the answer, I’ll answer it anyway.” Me too. I think I know the answer to all control questions. The only problem is, a lot of my answers are wrong.
I’ve learned to differentiate between wrong and right answers by trial and error. If the panel board operator persistently prefers to run a new control loop that I’ve designed in manual, if he switches out of auto whenever the flow becomes erratic, then I’ve designed a control strategy that’s wrong. So, that’s how I’ve learned to discriminate between a control loop that works and a control strategy best forgotten.
Here’s something else I’ve learned. Direct from Dr. Shinsky, the world’s expert on process control:
“Lieberman, if it won’t work in manual, it won’t work in auto.”
“Most control problems are really process problems.”
I’ve no formal training in process control and instrumentation. All I know is what Dr. Shinsky told me. And 59 years of experience in process plants has taught me that’s all I need to know.
LEARNING FROM PLANT OPERATORS
My first assignment as a Process Engineer was on No. 12 Pipe Still in Whiting, Indiana. This was a crude distillation unit. My objective was to maximize production of gas oil, as shown in Figure 1‐1. The gas oil had a product spec of not more than 500 ppm asphaltenes. The lab required half a day to report sample results. However, every hour or two, the outside operator brought in a bottle of gas oil for the panel board operator. The panel operator would adjust the wash oil flow, based on the color of the gas oil.
Figure 1‐1Adjusting wash oil based on gas oil color
While plant supervision monitored the lab asphaltene sample results, plant operators ignored this analysis. They adjusted the wash oil rate to obtain a clean‐looking product. The operators consistently produced a gas oil product with 50–200 ppm asphaltenes. They were using too much wash oil, and the more the wash oil used, the lower the gas oil production.
I mixed a few drops of crude tower bottoms in the gas oil to obtain a bottle of 500 ppm asphaltene material. I then instructed the panel board operators as follows:
If the sample from the field is darker than my standard bottle, increase the wash oil valve position by 5%.
If the sample of gas oil from the field is lighter than my standard, decrease the wash oil valve position by 3%.
Repeat the above every 30 minutes.
The color of gas oil from a crude distillation unit correlates nicely with asphaltene content. The gas oil, when free of entrained asphaltenes, is pale yellow. So, it seems that my procedure should have worked. But it didn’t. The operators persisted in drawing the sample every 1–2 hours, not every 30 minutes like I had instructed.
So, I purchased an online colorimeter. The online colorimeter checked whether the gas oil color was above or below my set point. With an interval of 10 minutes, it would move the wash oil valve position by 1%. This never achieved the desired color, but the gas oil product was mixed in a tank. The main result was that gas oil production was maximized, consistent with the 500 ppm asphaltene specification.
One might say that all I did was automate what the operators were already doing manually and that all I accomplished was marginally improving an existing control strategy by automating the strategy. But, in 1965, I was very proud of my accomplishments. I had proved, as Dr. Shinsky said, “If it does work on manual, we can automate it.”
LEARNING FROM FIELD OBSERVATIONS
Fifty‐three years ago, I redesigned the polypropylene plant in El Dorado, Arkansas. I had never paid much attention to control valves. I had never really observed how they operate. But I had my opportunity to do so when the polypropylene plant was restarted.
The problem was that the purchased propylene feed valve was too large for normal service. I had designed this flow for a maximum of 1600 BSD, but the current flow was only 100 BSD. Control valve response is quite nonlinear. Nonlinear means that if the valve is open by 5%, you might get 20% of the flow. If you open the valve from 80 to 100%, the flow goes up by an additional 2%. Nonlinear response also means that you cannot precisely control a flow if the valve is mostly closed. With the flow only 20% of the design flow, the purchased propylene feed was erratic. This resulted in erratic reactor temperature and erratic viscosity of the polypropylene product.
The plant start‐up had proceeded slowly. It was past midnight. The evening was hot, humid, and very dark. I went out to look at the propylene feed control valves. Most of the flow was coming from the refinery’s own propylene supply. This valve was half open. But the purchased propylene feed valve was barely open. The valve position indicator, as best I could see with my flashlight, was bumping up and down against the “C” (closed) on the valve stem indicator.
The purchased propylene charge pump had a spillback line, as shown in Figure 1‐2. I opened the spillback valve. The pump discharge pressure dropped, and the propylene feed valve opened to 30%. The control valve was now operating in its linear range.
Now, when I design a control valve to handle a large reduction in flow, I include an automated spillback valve from pump discharge to suction. The spillback controls the pump discharge pressure to keep the Flow Recorder Control (FRC) valve between 20 and 80% open. Whenever I sketch this control loop, I recall that dark night in El Dorado. I also recall the value of learning even the most basic control principles by personal field observations.
Figure 1‐2Opening spillback to keep the FRC valve in its linear operating range
LEARNING FROM MISTAKES
Adolf Hitler did not always learn from his mistakes. For example, he once ordered a submarine to attack the Esso Lago Refinery in Aruba. The sub surfaced in the island’s harbor and fired at the refinery. But the crew neglected to remove the sea cap on the gun’s muzzle. The gun exploded and killed the crew.
I too had my problems in this refinery. The refinery flare was often very large and always erratic. The gas being burned in the flare was plant fuel. The plant fuel was primarily cracked gas from the delayed coker, supplemented (as shown in Fig. 1‐3) by vaporized LPG. So much fuel gas was lost by flaring that 90% of Aruba’s LPG production had to be diverted to fuel, via a propane vaporizer, to maintain refinery fuel gas pressure.
I analyzed the problem based on the dynamics of the system. I modeled the refinery’s fuel consumption versus cracked gas production as a function of time. The key problem, based on my computer system dynamic analysis, was the cyclic production of cracked gas from the delayed coker complex. My report to Mr. English, the General Director of the Aruba Refinery, concluded:
The LPG vaporizer was responding too slowly to changes in cracked gas production from the delayed coker.
The natural log of the system time constants of the coker and vaporizer was out of synchronization.
A feed‐forward, advanced computer control based on real‐time dynamics would have to be developed to bring the coker vaporizer systems into dynamic real‐time equilibrium.
A team of outside consultants, experts in this technology, should be contracted to provide this computer technology.
Six months passed. The complex, feed‐forward computer system was integrated into the LPG makeup and flaring controls shown in Figure 1‐3. Adolf Hitler would have been more sympathetic than Mr. English. The refinery’s flaring continued just as before. Now what?
Figure 1‐3Unintentional flaring caused by malfunction of the LPG makeup control valve is an example of split ‐ range pressure control
Distressed, discouraged, and dismayed, I went out to look at the vaporizer. I looked at the vaporizer for many hours. After a while, I noticed that the fuel gas system pressure was dropping. This happened every 3 hours and was caused by the cyclic operation of the delayed coker. This was normal.
The falling fuel gas pressure caused the instrument air signal to the LPG makeup valve to increase. This was an “Air‐to‐Open” valve (see Chapter 13), and more air pressure was needed to open the propane flow control valve. This was normal.
But, the valve position itself did not move. The valve was stuck in a closed position. This was not normal.
You will understand that the operator in the control room was seeing the LPG propane makeup valve opening as the fuel gas pressure dropped. But the panel board operator was not really seeing the valve position; he was only seeing the instrument air signal to the valve.
Suddenly, the valve jerked open. The propane whistled through the valve. The local level indication in the vaporizer surged up, as did the fuel gas pressure. The flare valve opened to relieve the excess plant fuel gas pressure and remained open until the vaporizer liquid level sank back down, which took well over an hour. This all reminded me of the sticky side door to my garage in New Orleans.
I sprayed the control valve stem with WD‐40, stroked the valve up and down with air pressure a dozen times, and cleaned the stem until it glistened. The next time the delayed coker cycled, the flow of LPG slowly increased to catch the falling fuel gas pressure, but without overshooting the pressure set point and initiating flaring.
My mistake had been that I had assumed that the field instrumentation and control valves were working properly. I did not take into account the probability of a control valve malfunction. But, at least, I had learned from my mistake, which is more than you could say for Adolf Hitler.
LEARNING FROM THEORY
Northwestern University has an excellent postgraduate chemical engineering program. I know this because I was ejected from their faculty. I had been hired to present a course to their graduate engineers majoring in process control. My lecture began:
“Ladies and gentlemen, the thing you need to know about control theory is that if you try to get some place too fast, it’s hard to stop. Let’s look at Figure 1‐4. In particular, let’s talk about tuning the reflux drum level control valve.
Do I want to keep the level in the drum close to 50%, or doesn’t it matter? As long as the level doesn’t get high enough to entrain light naphtha into fuel gas, that’s okay. What is not okay is to have an erratic flow feeding the light naphtha debutanizer tower.
On the other hand, if the overhead product was flowing into a large feed surge drum, than precise level control of the reflux drum is acceptable.
In order for the instrument technician to tune the level control valve, you have to show him what you want. To do this, put the level valve on manual. Next, manipulate the light naphtha flow to permit the level swings in the reflux drum you are willing to tolerate. But you will find that there is a problem. If you try to get back to the 50% level set point quickly, you will badly overshoot your level target.
Figure 1‐4Tuning a level control valve depends on what is downstream
If you return slowly to the set point, it’s easy to reestablish the 50% level target. However, the level will be off the target for a long time.
In conclusion, ladies and gentlemen, tuning a control loop is a compromise between the speed at which we wish to return to the set point and our tolerance to overshooting the target. To establish the correct tuning criteria, the control loop is best run on manual for a few hours by the Process Control Engineer. Thank you. Class adjourned for today.”
My students unfortunately adjourned to Dean Gold’s office. Dean Gold lectured me about the student’s complaints.
“Mr. Lieberman, did you think you were teaching a junior high school science class or a postgraduate course in process control?”
And I said, “Oh! Is there a difference?”
So that’s how I came to be ejected from the faculty of Northwestern University after my first day of teaching.
LEARNING FROM RELATIONSHIPS
My ex‐girlfriend used to tell me, “Norm, the reason we get along so well is that I give you a lot of positive feedback.” From this, I developed the impression that positive feedback is good, which is true in a relationship with your girlfriend. But when involved in a relationship with a control loop, we want negative feedback. Control logic fails when in the positive feedback mode of control. For example:
Distillation
—As process engineers and operators, we have the expectation that reflux improves fractionation, which is true, up to a point. That point where more reflux hurts fractionation instead of helps is called the “incipient flood point.” Beyond this point, the distillation tower is operating in a positive feedback mode of process control. That means the tray flooding reduces tray fractionation efficiency. More reflux and more reboiler heat simply make the flooding worse.
Fired Heaters
—Increasing furnace fuel should increase the heater outlet temperature. But if the heat release is limited by combustion air, then increasing the fuel gas will reduce the heater outlet temperature. But as the heater outlet temperature drops, the automatic control calls for more fuel gas, which does not burn. As the heater outlet temperature continues to fall, because combustion is limited by air, the outlet temperature drops further. The heater automatic temperature control loop is now in the positive feedback mode of control. As long as this control loop is on auto, the problem will feed upon itself.
Vacuum Ejector
—Some refineries control vacuum tower pressure by controlling the motive steam flow to the steam ejector. As the steam pressure and flow to the ejector increases, the ejector pulls a better vacuum, as shown in
Figure 1‐5
, but as the steam flow increases, so does that load on the downstream condenser. As the condenser becomes overloaded, the ejector discharge pressure rises. At some point, the increased discharge pressure adversely affects the ejector’s suction pressure. A further increase in motive steam will make the vacuum worse, instead of better. As the vacuum gets worse, the control loop calls for more steam. Having now entered the positive feedback mode of control, the problem feeds upon itself.
Figure 1‐5Too much steam flow causes a loss in vacuum
Many control loops are subject to slipping into a positive feedback loop. The only way out of this trap is to switch the controls to manual and slowly climb back out of the trap. Once you guess (but there is no way to know for sure) that you are in the safe, negative feedback mode of control, you can then safely switch back to automatic control.
NORMAL PURPOSE OF CONTROL LOOPS
Typically, a control loop is tuned to achieve two objectives:
To return a variable to its set point as fast as possible.
To avoid overshooting the set point.
If a heater outlet set point is at 700°F, and it is currently running at 680°F, the firing rate should increase. However, if the firing rate increases too fast, the heater outlet may jump past the set point to 720°F.
Tuning a control loop is meant to balance the instrument, “gain and reset,” to balance these two contradictory objectives.
The balance between gain and reset (i.e., instrument tuning) is not the main object of this text. Only rarely have I seen a panel board operator complain about this problem.
Another purpose of control is to optimize process variables. This is an advanced control that attempts to optimize certain variables. This is also not the sort of problem that the panel operator would be concerned about. An example of advanced control would be to optimize the ratios of several pumparounds, versus the top reflux rate, for a refinery crude distillation tower. For the units I work on, such advanced computer control is rarely used, or has been simplified, so that it is not much different than ordinary closed‐loop control.
MANUAL VERSUS AUTO
In reality, the main complaint about control loops that are communicated to me by operators is that the controller will not work in the automatic mode of control and that the operators are forced to run the control loop in manual. This greatly increases and complicates their work.
To a large extent, this text examines why control loops are forced to run in the manual mode. A few of the reasons are the following:
The control loop is trapped in a “positive feedback loop.” This is often a dangerous situation.
There is no direct relationship between the variable being controlled and the response of the control valve. This is typically a design error.
The facility that measures the process parameter in the field is not working correctly. This represents the majority of control problems that I have seen.
The bypass valve is open around the control valve.
The control valve is running too far closed because it is oversized or badly eroded.
The control valve is running too far open because it is too small, or its port size is too small, or an isolation gate valve in the system is partly closed.
The control valve’s “Hand Jack” has been left engaged. Thus, the control valve cannot be manipulated from the computer console or panel. The hand jack is a mechanical device, used to manually move the control valve in the field.
The control valve is stuck in a fixed position.
The air signal connection to the diaphragm that moves the control valve has come loose.
The diaphragm is leaking, so the sufficient instrument air pressure cannot be applied to the control valve mechanism to force it to move.
The air signal is connected on the wrong side of the diaphragm.
PROCESS CONTROL NOMENCLATURE
The reader who is new to process plant vocabulary may wish to briefly skip to the glossary at the end of this book. I have assembled a list of “Process Control Nomenclature Used in Petroleum Refineries and Petrochemical Plants.” As in any other industry, your coworkers will have developed a vocabulary of their own and will assume you understand the terms they employ. To an extent, in the following chapters of this text, I have also made a similar assumption.
A brief review of these terms may make it easier for you to communicate with some of your coworkers.
