industrial gas handbook gas separation and purification
ß 2006 by Taylor & Francis Group, LLC.
ß 2006 by Taylor & Francis Group, LLC.
industrial gas handbook gas separation and purification
frank g. Kerry
ß 2006 by Taylor & Francis Group, LLC.
CRC Press Taylor & Francis Group 6000 Broken Sound Parkway NW, Suite 300 Boca Raton, FL 33487-2742 © 2007 by Taylor & Francis Group, LLC CRC Press is an imprint of Taylor & Francis Group, an Informa business No claim to original U.S. Government works Printed in the United States of America on acid-free paper 10 9 8 7 6 5 4 3 2 1 International Standard Book Number-10: 0-8493-9005-2 (Hardcover) International Standard Book Number-13: 978-0-8493-9005-0 (Hardcover) This book contains information obtained from authentic and highly regarded sources. Reprinted material is quoted with permission, and sources are indicated. A wide variety of references are listed. Reasonable efforts have been made to publish reliable data and information, but the author and the publisher cannot assume responsibility for the validity of all materials or for the consequences of their use. No part of this book may be reprinted, reproduced, transmitted, or utilized in any form by any electronic, mechanical, or other means, now known or hereafter invented, including photocopying, microfilming, and recording, or in any information storage or retrieval system, without written permission from the publishers. For permission to photocopy or use material electronically from this work, please access www. copyright.com (http://www.copyright.com/) or contact the Copyright Clearance Center, Inc. (CCC) 222 Rosewood Drive, Danvers, MA 01923, 978-750-8400. CCC is a not-for-profit organization that provides licenses and registration for a variety of users. For organizations that have been granted a photocopy license by the CCC, a separate system of payment has been arranged. Trademark Notice: Product or corporate names may be trademarks or registered trademarks, and are used only for identification and explanation without intent to infringe. Library of Congress Cataloging-in-Publication Data Kerry, Frank G. Industrial gas handbook : gas separation and purification / Frank G. Kerry. p. cm. Includes bibliographical references and index. ISBN 0-8493-9005-2 1. Gases--Separation--Handbooks, manuals, etc. 2. Gases--Purification--Handbooks, manuals, etc. I. Title. TP242.K47 2006 665.7--dc22 Visit the Taylor & Francis Web site at http://www.taylorandfrancis.com and the CRC Press Web site at http://www.crcpress.com
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2006049163
Memoriam In Memory of my wife Mary Jane who patiently undertook the role of Penelope while I emulated the ever-roving Odysseus in search of new gases to recover and purify.
ß 2006 by Taylor & Francis Group, LLC.
ß 2006 by Taylor & Francis Group, LLC.
Preface As a practicing engineer in the field of industrial gases for the past 65 years, 40 with L’Air Liquide, and another 25 as a private consultant on an international basis, I have encountered only seven textbooks worldwide compiled for our industry, even though the industry is over 100 years old. Of the seven, five were edited by university professors or academic research scientists. The sixth was written in 1989 by a process manager of an industrial company in England which almost filled the bill, but not quite. Those were published in England, one in France, two in the United States, and one in Russia. With the possible exception of the one on process cycles published in England, the rest were heavy on the theoretical side and did not serve the majority of the engineers’ need for applied technology. The purpose of this manuscript, based on many years of field experience, is not so much for physicists or scientists who occasionally may require data relating to cryogenics and gas recovery, but for practicing engineers to help them master the fundamentals and advances in the field of industrial gas separation and purification. In planning the breadth of this manuscript the following colleagues were most helpful: Ted Pawulski. A former director of process design for separation units for Air Liquide in North America, now a private consultant, who reviewed the technical aspects of the text with diligence, patience, and made important suggestions. Nat Matlin. A graduate chemist and a polished copywriter by profession who went over the manuscript with a fine tooth comb. He was able to simplify the presentation and ease the flow of reading. To that end, he eliminated much of the technobabble so prevalent in current engineering reports. Joe Bernstein. A graduate chemical engineer with a stint at Air Products & Chemicals, now a private consultant, reviewed Chapter 3 on distillation and made useful recommendations. Homage must also be paid to the countless engineer-technologists whom I met in my professional life, who spend a lifetime in the field and away from their families, supervising the erection of machinery and equipment, the installation of instrumentation, the burdensome stroking of control valves, the startup of process units, and writing reports to head office, at the same time mumbling audibly that this is the last project they will participate in. And yet, the same faces show up in other projects, other countries, or other continents. They have amassed a precious lode of practical experience, which should be mined by enterprising young engineers at every opportunity. Many thanks should also go to my daughter, Nicolette Kerry, who functioned as a full secretarial staff correcting typos, maintaining my computer, running off copies, contacting outside sources for data and serving as both my most severe critic and most ardent supporter. This book is indebted to Elizabeth M. Mihaltse, whose professional graphic design expertise made a difference for the better, in many of the figures found in her grandfather’s book. Aside from individuals, the following organizations, listed alphabetically, who offered beneficial advice, as well as granting permission for use of data, diagrams, and photographs should be thanked: . . .
A.I.Ch.E Air Liquide Process & Construction American Institute of Physics
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. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
ASTM International Atlas Copco BOC. Group BP Corp. Burckhardt Compression Carbo Tech Engineering GmbH Chart Energy & Chemicals, Inc. Chemical Engineering Chemical Engineering Progress (CEP) Compressed Gas Manufacturing Association Crane Co. Cryogenic Consulting Service, Inc. CryoGas International Consolidated Design Descon Engineering Co. Inc. ExxonMobil Upstream Affairs G.A.O. Hydrocarbon Processing (Gulf Publishing) Ingersol Rand Company Limited Kluwer Academic KOSO America Inc. (Hammel -Dahl) Leybold Vacuum LindeBOC Process Plants Los Angeles County Sanitation Districts MANTurbo AG Burckhardt Compressors McGraw Hill Education MV Engineering Chart Industries Metso Automation Oil & Gas Report Oxford University Press Perlite Institute Plenum Press Praxair, Inc. Southern Research Institute (University of Alabama) Bates & Hussman Springer Science & Business Media Taylor Wharton Thermax Thomson Learning Global Right Group Tyco Valves & Controls W.H. Freeman & Company
Nota Bene The reader is advised that because international industry uses SI units we have also used these with the possible exception of temperatures above 273 K where Celsius may be more convenient, especially for metallurgists; also US gallons for pumps and barrels for crude oil. To convert 8F into K, one may use K ¼ 5=9 (F þ 459.7) for all cases including if F is minus. Although kPa is used for small pressures up to and slightly above atmospheric, the term bar for higher pressures is preferable because the majority of international compressor
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manufacturers employ it. In fact, if one deducts the pressure drop of an air filter of an air separation unit the process air enters the compressor at approximately 1 bar. 101.325 kPa ¼ 1 atm (most engineers use 101.3 kPa) 1 bar ¼ 100 kPa The term heat gain is used because heat enters the cryogenic equipment increasing the temperature and does not leak out. For vacuum applications the norm is bar, mbar, mbar, and Pascal.
ß 2006 by Taylor & Francis Group, LLC.
ß 2006 by Taylor & Francis Group, LLC.
Author Frank George Kerry is President of F.G. Kerry, Inc. with offices in Scarsdale, N.Y. and Sao Paulo, Brazil. He has been a consultant in the field of cryogenic systems and industrial gas technology for the past 30 years. He has directed major gas and air separation projects on six continents including the Dakota Gasification Company (USA), Companhia Siderurgica Paulista (Brazil), Siderurgica del Orinoco (Venezuela), and Barrick Goldstrike Mines (USA). Before establishing his own consulting organization, he spent 40 years with L’Air Liquide Group, first in Canada, then in New York where he established L’Air Liquide’s US operations. While being appointed vice-president, and serving as a member of the board he initiated international corporate development studies for the world headquarter in Paris, France. An honor graduate and medalist of McGill University (B. Eng Civil), Mr. Kerry continued on to do postgraduate studies in Paris, France (Ing. ESSA). During World War II, he was on loan to the Canadian Government to establish welding procedures for the fabrication of maritime tankers. He also designed and supervised the fabrication of piping networks to distribute oxygen and acetylene in shipyards, airplane factories, and steel mills. After the war, he concentrated on the application of oxygen in the open hearth furnace, and in 1946 he was granted a US patent (no. 2,446,511). He has presented numerous lectures in the USA and abroad, and authored many articles especially on gas purification and safety. He is a Member Emeritus of A.I.Ch.E, ASTM International, and ASM (Brazil).
ß 2006 by Taylor & Francis Group, LLC.
ß 2006 by Taylor & Francis Group, LLC.
Table of Contents Chapter 1 Gas Separation and Purification of Industrial Gases 1.1 Introduction 1.2 Thermodynamics 1.2.1 General Principles of Thermodynamics 1.2.2 Enthalpy (H) (J=kg) 1.2.3 The Second Law (The Availability of Energy) 1.2.4 Carnot Cycle 1.2.5 Entropy (S ) 1.2.5.1 Irreversible Systems 1.2.6 Third Law 1.2.7 Real Gases 1.2.8 Compression of Gases 1.2.8.1 Critical Temperatures and Pressures 1.2.9 Compressibility 1.2.10 Free Expansion through a Valve 1.2.11 Inversion 1.2.11.1 Deviation from Boyle’s Law 1.2.12 Adiabatic Expansion 1.2.13 Thermodynamic Charts and Tables 1.2.14 Cryogenic Properties of Air 1.2.15 Refrigeration and Liquefaction Systems (Ideal and Reversible) 1.2.16 Vapor Compression Systems 1.2.17 Liquefaction Systems 1.2.17.1 High-Pressure Free Expansion (Isenthalpic Linde–Hampson) System 1.2.17.2 Claude Isentropic System 1.2.17.3 Precooling Systems 1.2.17.4 Cascade Systems 1.2.18 Summary References Chapter 2 Industrial Applications 2.1 Early Development of Industrial Liquefaction Systems 2.2 Heat Exchangers 2.3 Expansion Machines 2.4 Contemporary Liquefaction Cycles 2.5 Linde Cycle (Free Expansion through a Valve) 2.5.1 Theoretical Analysis of the First Linde High Pressure Cycle 2.5.2 Theoretical Analysis of Linde Basic Cycle with Precooling 2.5.3 Theoretical Analysis of the Linde High-Pressure Dual Process
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2.6
Theoretical Analysis of the Claude Cycle 2.6.1 Claude Cycle with Precooling 2.6.2 Claude Cycle with Dual Pressures 2.6.3 Claude Cycle with High Precooling to Liquefy Hydrogen or Neon 2.6.4 The Low-Temperature Refrigerator 2.7 Kapitza Cycle 2.8 Cascade Cycle References Further Reading Chapter 3 Air Separation Technology 3.1 Air Separation Overview 3.1.1 Linde’s First Fractionation Machine 3.1.2 Distillation and Fractionation 3.1.3 Fractionation 3.1.4 Stripping 3.1.5 Rectification 3.2 Theoretical Considerations of Fractionation 3.2.1 Evaporation and Condensation 3.2.2 Simple Separation by Condensation and Flashing (Separators) 3.2.2.1 Application 3.2.2.2 Procedure 3.2.2.3 Practical Example (It Requires Iteration) 3.2.3 Fractionation 3.2.4 Fractionation Methods 3.2.5 Fractionation Plates 3.2.6 Analysis of Flow in Equipment 3.2.6.1 Case I: Analysis of a Low Pressure Column 3.2.6.2 Case II: Consideration of Vapor Feed 3.2.6.3 Case III: Analysis of the High-Pressure Column 3.3 Practical Considerations 3.3.1 Bubble-Cap Trays 3.3.2 Sieve Trays 3.3.3 Structured Packings 3.3.4 Care in the Design of Structured Packing 3.3.5 Safety in the Use of Structured Packing 3.4 Operational Control 3.4.1 FCV-1 3.4.2 FCV-2 3.4.3 FCV-3 3.4.4 Refrigeration 3.5 Product Recovery 3.6 Optimum Reflux 3.7 Distillation Equipment 3.7.1 Upper (Low Pressure) Column 3.7.2 Lower (High Pressure) Column 3.7.3 Main Condenser 3.7.4 Liquid Subcoolers
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3.7.5 Process Considerations 3.7.6 Crude-Argon Separation Column 3.8 Development of Low Oxygen-Purity Processes 3.8.1 The Lachmann Principle 3.8.2 The Oxyton Process 3.8.2.1 Thermodynamic Analysis of the Oxyton Cycle 3.8.2.2 Oxyton Development 3.8.3 Variable-Load Plants 3.8.3.1 Version A 3.8.3.2 Version B 3.9 Exergy References Chapter 4 Rare (Noble) Gases 4.1 Helium 4.1.1 Sources of Helium 4.1.2 General Principles of Recovery of Helium 4.1.3 Recovery Processes from Natural Gases 4.1.4 Applications of Helium 4.1.5 Conservation of Helium 4.2 Neon 4.2.1 General 4.2.2 Sources of Neon 4.2.3 Recovery of Neon 4.2.4 Industrial Recovery of Neon 4.2.5 Industrial Applications of Neon 4.3 Argon 4.3.1 General 4.3.2 Sources of Argon 4.3.3 Recovery of Argon 4.3.4 Recovery Procedure and Equipment 4.3.5 Secondary Rectification and Final Purification 4.3.6 Refining Operation and Equipment 4.3.7 Applications of Argon 4.4 Krypton and Xenon 4.4.1 General 4.4.2 Sources of Krypton and Xenon 4.4.3 Recovery of Krypton and Xenon 4.4.4 Refining of Krypton and Xenon 4.4.5 Recovery of Rare Gases from Ammonia Purge Gas 4.4.6 Applications of Krypton and Xenon References Further Reading Chapter 5 Front-End Purification Systems 5.1 Historical Background 5.1.1 Processes and Materials Used in Front-End Purification Systems
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5.1.2 5.1.3
Original Prepurification Adsorbents 5.1.3.1 General 5.1.4 Introduction of Activated Alumina 5.1.4.1 Activated Alumina 5.1.4.2 Regeneration of Activated Alumina 5.1.5 Zeolites (Molecular Sieves) 5.1.5.1 Chemical Formula 5.1.5.2 Types of Molecular Sieves 5.1.6 Silica Gel 5.2 Design of Current Front-End Purification Systems 5.2.1 General Background 5.2.2 Equipment Used 5.2.2.1 Precooling Units Upstream of Adsorption 5.2.2.2 Direct Contact Aftercooler 5.2.2.3 Evaporative Water Chiller 5.2.2.4 Mechanical Refrigeration Unit 5.2.3 Adsorber Unit 5.2.3.1 Standard Design 5.2.3.2 Multiple Vertical Vessels 5.2.3.3 Horizontal Vessels 5.2.3.4 Radial or Concentric Design 5.3 Process Operation 5.3.1 Isolation Valve Downstream of FEP 5.3.2 Adsorption Kinetics 5.3.3 Regeneration Concerns 5.3.4 Warning against Excessive Heat during Regeneration 5.3.5 High-Pressure Vessel Regeneration 5.3.6 Regeneration Options for FEP Units 5.3.6.1 General 5.3.7 Summary 5.3.8 Operational Time Cycle 5.3.9 Prepurification Adsorbent Units and Operating Stability 5.3.9.1 Improving Operating Stability 5.4 Safety 5.4.1 Hydrocarbon Breakthrough 5.4.2 Safety Add-ons 5.4.3 Liquid Oxygen Purge 5.4.4 Analyzers 5.5 Activated Aluminas for Front-End Purification Systems 5.5.1 Background 5.5.2 Pressure Swing Adsorption 5.5.3 Industrial Applications in Air Separation Plants 5.5.4 Observations on PSA Prepurification 5.5.5 Field Observations References Further Reading on the Subject of Adsorption and Carbon Dioxide Build-up
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Chapter 6 Product Liquefaction, Storage, and Transportation 6.1 Background 6.1.1 New Applications 6.2 Product Liquefaction 6.2.1 Enthalpy Balance 6.2.2 Direct Extraction 6.2.3 Basic Design Parameters for an Efficient Liquefaction System 6.2.4 Types of Liquefiers 6.2.4.1 Independent Liquefier 6.2.4.2 Integrated Liquefier 6.2.4.3 Very High-Pressure Liquefiers 6.2.4.4 General Summary 6.2.4.5 Energy Economics 6.3 Cryogenic Storage Facilities 6.3.1 General Considerations 6.3.2 Geographic Considerations 6.3.2.1 Ambient Temperature 6.3.2.2 Wind Loading 6.3.2.3 Seismic Loadings 6.3.2.4 Soil Conditions and Land Cost 6.3.2.5 Snow Loads 6.3.2.6 External Corrosion 6.3.2.7 Availability and Dependability of Utilities 6.3.2.8 Local Neighborhood Characteristics 6.3.3 Design Parameters 6.3.3.1 Low-Pressure Shop-Built Tanks 6.3.3.2 Storage Vessels with Internal Pressure 6.3.3.3 Low-Pressure Field-Built Aluminum Tanks 6.3.3.4 Flat Bottom Tanks 6.3.3.5 Spherical Containers 6.3.3.6 Cylindrical Vessels (Horizontal or Vertical) 6.3.4 Design Selection 6.3.5 Typical Designs of Cryogenic Storage Vessels 6.3.5.1 Vertical Cylindrical Tanks 6.3.5.2 Horizontal Cylindrical Storage Tanks 6.3.5.3 Spherical Tanks 6.3.5.4 Flat Bottom Tanks 6.3.6 Cryogenic Liquid Delivery Systems 6.3.6.1 General 6.3.6.2 Small Portable Containers 6.3.6.3 Customer Bulk Stations 6.3.6.4 LOX Distribution in a Shop 6.3.6.5 Liquid Deliveries by Truck 6.4 Cryogenic Pumps 6.4.1 Background 6.4.2 Variety of Applications 6.4.3 Materials 6.4.4 Present Designs 6.4.5 Net Positive Suction Head
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6.4.6 6.4.7 6.4.8
Inlet Filter Screen Installation and Connections Typical Pump Piping Calculations 6.4.8.1 Velocity Constraints of Cryogenic Fluids 6.4.8.2 Pressure Drops due to Piping Components 6.4.9 Start-up of Pumps 6.4.10 High-Pressure Radial Pumps 6.4.11 Ultrahigh Pressure Pumps 6.4.12 Automation 6.5 Cryogenic Liquid Vaporizers 6.5.1 General Overview 6.5.2 Ambient Air Vaporizers 6.5.2.1 Modules Spread Apart 6.5.2.2 Modules in Alternate Operation 6.5.2.3 Modules with Pressurized Air 6.5.3 Direct Steam Vaporizers 6.5.3.1 Vaporization with Steam-Heated Water 6.5.4 Emergency Vaporization of Products References Chapter 7 Insulation 7.1 General 7.1.1 Theoretical Considerations 7.1.2 Insulations: General 7.1.3 Vacuum Insulation (Radiation) 7.1.4 Conductivity in Mass Insulations 7.1.5 Natural Convection in Mass Insulation 7.1.6 Vacuum Plus Powder or with Fibrous Insulations 7.1.7 Insulation (Multilayer, Super, or Simply MLI) 7.2 Industrial Practices 7.2.1 Industrial Applications of Insulation 7.2.2 Cryogenic Casings (Cold Boxes) for Process Equipment 7.2.3 Mineral Wool (Rock Wool) 7.2.4 Expanded Perlite 7.2.5 Glass Wool (Fiberglass) 7.2.6 Glass Blocks (Foam Glass) 7.2.7 Vermiculite 7.2.8 Silica Aerogel 7.2.9 Magnesium Carbonate 7.3 Cold Box Design for Insulation 7.3.1 Special Requirements for Liquid Hydrogen Processing Plants 7.4 Externally Located Process and Transfer Piping 7.4.1 Short Lines 7.4.2 Expanded Foams 7.4.3 Fiberglass Insulation 7.4.4 Prefabricated Vacuum-Insulated Piping 7.4.5 Multilayer Insulation 7.4.6 Cryogenic Liquid Piping Design
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7.5
Insulation for Liquid Storage Tanks and Vessels 7.5.1 Large Storage Tanks (1000 t and over) 7.5.2 Smaller Storage Tanks (500 to 1000 t or 500 to 1000 kL) 7.5.3 Storage Tanks (50 to 500 t or 50 to 500 kL) 7.5.4 Storage Vessels (up to 50 t or 50 kL) 7.6 Vacuum Pumping Systems 7.6.1 General Overview 7.6.2 Vacuum Pumps 7.6.2.1 Roots Vacuum Pump 7.6.2.2 Rotary Vacuum Pump 7.6.2.3 Turbomolecular Pumps 7.6.2.4 Cryopumps 7.6.2.5 Adsorption Pumps 7.6.2.6 Getters 7.6.2.7 Small Laboratory Pumps 7.6.3 Periodic Purging and Deriming 7.6.4 Ancillary Equipment 7.6.4.1 Valves 7.6.4.2 Vacuum Measurement References For Further Study and Review Chapter 8 Special Gases 8.1 Hydrogen 8.1.1 Sources of Hydrogen 8.1.2 Recovery of Hydrogen 8.1.3 Hydrogen Use in Petroleum Refineries 8.1.4 Refinery In-House Recovery of Hydrogen 8.1.5 Recovery from Coke Oven Gas 8.1.6 Hydrogen Generation Plants 8.1.6.1 Electrolysis of Water 8.1.6.2 Thermal Cracking of Ammonia 8.1.6.3 Treatment of Hydrocarbon Feedstock for Hydrogen Recovery 8.1.6.4 Small Steam Reforming Plants (150–1000 Nm3=h) 8.1.6.5 Large Hydrogen Generation Plants (over 1000 Nm3=h) 8.1.7 Synthesis Gas, Partial Oxidation 8.1.7.1 History 8.1.7.2 Partial Oxidation Process 8.1.7.3 Ammonia Synthesis 8.1.7.4 Hydrogen Recovery from Ammonia Synthesis Plants 8.1.7.5 Other Uses for Synthesis Gas 8.1.7.6 Fuel Cells 8.2 Carbon Monoxide 8.2.1 Sources 8.2.2 Carbon Monoxide Recovery 8.2.2.1 General 8.2.3 General Process of Recovery 8.2.4 Basic Cryogenic Recovery Processes
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8.3
8.4
8.5
8.6
8.2.4.1 Methane Wash Cryogenic Recovery 8.2.4.2 Simplified Carbon Monoxide Recovery 8.2.5 Compression and Conversion Machinery for Carbon Monoxide 8.2.5.1 Requirements for a Liquid CO Pump 8.2.5.2 Conversion from LCO to Gas 8.2.5.3 Safety of LCO Transport 8.2.6 Safety Equipment in General High-Purity Nitrogen 8.3.1 General Characteristics 8.3.2 Recovery 8.3.3 Applications for Inertness 8.3.3.1 Nitrogen as a Preservative 8.3.3.2 Nitrogen as an Emissions Controller 8.3.3.3 Nitrogen Use in Sparging 8.3.3.4 Nitrogen for Pressure Transferring 8.3.3.5 Liquid Nitrogen for Vapor Recovery 8.3.3.6 Liquid Nitrogen Makes Worn Rubber Tires Profitable 8.3.4 Process and Equipment Options 8.3.4.1 Cryogenic Process Cycle 8.3.4.2 Permeable Membrane Separation Process 8.3.4.3 Pressure Swing Adsorption 8.3.5 Ultrahigh-Purity Nitrogen 8.3.5.1 Removal of Outside Impurities 8.3.5.2 Process Cycle for Ultrahigh Purity Nitrogen 8.3.5.3 Outside Factors in Contamination 8.3.6 Other Atmospheric Nitrogen Compounds 8.3.6.1 General 8.3.6.2 Dinitrogen Monoxide (N2O) or Laughing Gas 8.3.6.3 Applications of Nitrous Oxide 8.3.6.4 Dangerous Side of Nitrous Oxide Carbon Dioxide 8.4.1 General Characteristics 8.4.2 Sources of Carbon Dioxide 8.4.3 Recovery Processes for Carbon Dioxide 8.4.3.1 Food Grade Recovery from Petroleum Off-Gases. 8.4.3.2 Food Grade Recovery from a Fermentation Source 8.4.3.3 Nonfood Grade Carbon Dioxide 8.4.4 Dry Ice: Food Grade 8.4.4.1 Production of Dry Ice 8.4.5 Applications of Carbon Dioxide Ozone 8.5.1 General 8.5.2 Properties of Ozone 8.5.3 Atmospheric Ozone Layer 8.5.4 Generation of Ozone 8.5.5 Applications Methane 8.6.1 Properties of Methane 8.6.2 High-Purity Methane for Chemicals 8.6.3 Natural Gas Peak Load Shaving
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8.6.4
8.6.5 8.6.6
8.6.7
8.6.8
Base Load Natural Gas Plants (LNG) 8.6.4.1 Cascade Cycle with Mixed Refrigerants in General 8.6.4.2 ARC Process Cycle 8.6.4.3 Further Development of Mixed Refrigerant Cycles 8.6.4.4 Heat Exchangers 8.6.4.5 Propane Refrigeration System Pritko Process Cycle 8.6.5.1 General Process Cycle Final Product Purification 8.6.6.1 Nitrogen Rejection 8.6.6.2 Helium Recovery Natural Gas Prepurification 8.6.7.1 Acid Gases (CO2, H2S) 8.6.7.2 Water Removal (2H2O) 8.6.7.3 Mercury Contamination 8.6.7.4 Mercaptans 8.6.7.5 Butane 8.6.7.6 Propane and Ethane Economics 8.6.8.1 LNG Economics Safety
8.6.9 References Further Reading on LNG
Chapter 9 Noncryogenic Separations 9.1 Permeable Membrane Separation 9.1.1 General Principles 9.1.2 Mechanical Design of Membranes 9.1.3 General Applications 9.1.3.1 Nitrogen Separation 9.1.3.2 Disadvantages of Membrane Separation 9.1.3.3 Hydrogen Recovery 9.2 Gas Separation by Adsorption 9.2.1 General Overview 9.2.1.1 Adsorption Processes Studies 9.2.1.2 Regeneration of Adsorbent 9.2.1.3 Hydrogen Recovery from Coke Oven Gas 9.3 Nitrogen Recovery 9.3.1 Carbon Adsorbent (Carbon Molecular Sieve) (CMS) 9.3.2 High-Purity Hydrogen Recovery 9.3.3 Oxygen Separation and Vacuum Pressure Swing Adsorption 9.3.3.1 Process Description 9.3.4 Engineering Design 9.3.4.1 Basic Principles 9.3.4.2 Disadvantages of Adsorption 9.3.4.3 Economics References Additional Reading on Noncryogenic Separations
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Chapter 10 10.1 Cryogenic Equipment, Materials, and Machinery 10.1.1 Heat Exchangers 10.1.1.1 General 10.1.2 Parameters of Design 10.1.3 Basic Principles 10.1.4 Typical Example for Designing Tubular Heat Exchangers 10.1.5 Brazed Aluminum Heat Exchangers 10.1.6 Effectiveness («) 10.1.7 Operability 10.1.8 Efficiency (h) 10.1.9 Industrial Applications 10.1.10 Development of Brazed Aluminum Heat Exchangers 10.1.10.1 Pressure Limitations 10.1.11 Vacuum Brazed Heat Exchangers 10.1.12 Mechanical Construction 10.1.13 Limitations 10.1.14 Operation and Maintenance References Further Reading 10.2 Expansion Machines 10.2.1 Expansion Machines 10.2.1.1 General 10.2.2 Reciprocating Expansion Engine 10.2.3 Radial Expansion Machines 10.2.4 Process Applications 10.2.5 Operational Factor (Air Separation Plants) 10.2.6 Refrigeration Availability 10.2.7 Process Technology 10.2.8 Expansion Turbine Efficiency 10.2.9 Expansion Turbine Losses 10.2.10 Measuring Efficiency 10.2.11 Various Expansion Turbine Systems 10.2.12 Mechanical Design Parameters 10.2.12.1 General 10.2.12.2 Operational Control 10.2.12.3 Shaft Speed (rpm) 10.2.12.4 Impeller Design 10.2.12.5 Materials of Construction 10.2.12.6 Bearings 10.2.12.6.1 Lubrication System 10.2.13 Instrumentation and Control 10.2.13.1 Process Control 10.2.13.2 Instruments Required 10.2.13.3 Computer Control (DCS) 10.2.14 Spares 10.2.15 General Applications for Expansion Machines References Supplementary Reading
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10.3
Compressors 10.3.1 Compressors 10.3.1.1 General 10.3.2 Definitions 10.3.3 Centrifugal Compressors 10.3.3.1 General Parameters of Design (Per Stage) 10.3.3.2 Speed Variations 10.3.3.3 Surge Limitations and Pumping 10.3.3.4 Effect of Moisture 10.3.3.5 Effect of Altitude 10.3.3.6 Compressor Ratio Changes 10.3.3.7 Multistaging 10.3.3.8 Cooling Effect 10.3.3.9 Specific Speed 10.3.3.10 Stonewalling 10.3.3.11 Bearings 10.3.3.12 Seals 10.3.3.13 Lubrication System 10.3.3.14 Inlet Guide Vanes 10.3.3.15 Diffusers 10.3.4 Axial–Centrifugal Compressors 10.3.4.1 Bearings and Seals 10.3.4.2 Lubrication System 10.3.4.3 Inlet Guide Vanes 10.3.5 Axial Compressors 10.3.6 Integrally Geared Centrifugal Compressors (API-Standard-672) 10.3.6.1 General Overview 10.3.6.2 Functional Components and their Design 10.3.6.2.1 Gas Side 10.3.6.2.2 Mechanical Power Side 10.3.6.2.3 Bearings 10.3.6.2.4 Seals 10.3.6.3 Economics 10.3.7 Product Oxygen Compressors (RIO) 10.3.7.1 General Overview 10.3.7.2 General Safety Parameters 10.3.7.3 Safe Operation of Centrifugal Oxygen Compressors 10.3.7.4 Ultrahigh-Pressure Oxygen Compressors 10.3.7.4.1 Summary 10.3.8 High-Pressure Labyrinth Piston Compressor 10.3.8.1 Labyrinth Piston Compressor 10.3.8.1.1 General 10.3.8.2 Basic Design for Achieving Oil-Free Operation 10.3.8.3 Sealing Systems 10.3.8.4 Labyrinths 10.3.8.5 Internal Operating Elements 10.3.9 Compressor Drivers 10.3.9.1 Motor Torque 10.3.9.2 Enclosures 10.3.9.3 Power Factor
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10.3.10
Operating Reliability versus Capital Costs of Compressors 10.3.10.1 Recommendations 10.3.11 Applicable Compressor Correction Factors Variables Summary of Factors Application of Factors References 10.4 Valves and Valve Stations for the Cryogenic Industry 10.4.1 General Design and Materials 10.4.2 General Design in Sizing 10.4.3 Sizing Parameters 10.4.4 Valve Categories 10.4.5 Nonmetallic Material 10.4.6 Manufacturers of Fluorinated Polymers 10.4.7 Warm End Switching Valves 10.4.7.1 Warm End Reversing Valves 10.4.7.2 Warm End Switching Valves for PSA Pre-Purification Systems 10.4.8 Flow Control Check Valves 10.4.9 Cryogenic Process Valves (General) 10.4.10 Hand-Operated Cryogenic Valves 10.4.11 Process Control Valves 10.4.12 Product Flow Control Valves 10.4.12.1 Gaseous Products 10.4.12.2 Liquid Products 10.4.13 Valve Connections 10.4.14 Insulation and Casing Designs for Cryogenic Valves 10.4.15 Liquid Purge Valves 10.4.16 Automatic Control of Cryogenic Valves 10.4.17 Cryogenic Liquid Storage Valves 10.4.18 Pressure Safety Relief Valves: Overview 10.4.18.1 Sizing for Pressure Safety Relief Valves (International Units) SI 10.4.18.2 Pilot-Operated Safety Valves 10.4.18.3 Pressure and Vacuum Relief Valves 10.4.18.4 Bursting Disks 10.4.18.5 Check Valves 10.4.19 Maintenance of Cryogenic Valves 10.4.20 Valve Stations: General 10.4.21 Valve Station Design 10.4.22 Destruction of a Pressure Reduction Station 10.4.22.1 Hypothetical Conclusions 10.4.23 Recommendations Applicable to Pressure-Reducing Stations Further Reading Chapter 11 Instrumentation and Controls 11.1 Overview 11.2 General Requirements
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11.3 11.4
Controls and Control System Philosophy Minimum Instrumentation 11.4.1 General 11.4.2 Air Filter 11.4.3 Air Compressor 11.4.4 Direct Contact Aftercooler (If Applicable) 11.4.5 Front-End Purification 11.4.6 Air Separation Unit 11.4.7 Oxygen Product Compressor 11.4.8 Nitrogen Product Compressor 11.4.9 Liquid Oxygen Storage Tank 11.4.10 Liquid Nitrogen or Liquid Argon Storage Tank 11.4.11 Cooling Water System 11.4.12 Lube Oil System 11.4.13 Alarms, Shutdowns, and Interlocks 11.4.14 Analyzers 11.5 Possible Specific Requirements of Owner or Operator 11.5.1 Scope 11.5.2 Codes and Standards 11.5.3 Operational Philosophy 11.5.4 Distributed Control System 11.5.5 Field Instruments 11.5.5.1 Level Instruments 11.5.5.2 Temperature Instruments 11.5.5.3 Flow Instruments 11.5.5.4 Valves 11.5.5.5 Transmitters 11.5.5.6 Vibration Instruments 11.5.5.7 Local Controllers 11.5.5.8 Pressure Instruments 11.5.5.9 General 11.5.6 Interconnections 11.5.6.1 Pre-Packaging 11.5.6.2 Large Transformers Chapter 12 Safety 12.1 Safety Overview 12.2 Chemistry of Ignition, Combustion, and Explosion 12.2.1 Source of Combustibles 12.2.2 Ignition Energy 12.3 Critical Areas in an Air Separation Plant 12.3.1 General Description 12.4 Purification Systems 12.4.1 Adsorption Systems 12.4.2 Reversing Heat Exchangers: Revex 12.4.3 Nonreversing Heat Exchangers (Primary Heat Exchangers) 12.4.4 Distillation Column and Main Condenser 12.4.5 Auxiliary Vaporizers
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12.4.6
12.5 12.6
12.7
12.8
12.9 12.10 12.11 12.12 12.13 12.14
12.15
Ancillary Equipment for Safety 12.4.6.1 Rich Liquid Filters 12.4.6.2 LOX Guard Filter 12.4.7 Liquid Oxygen Storage Tanks 12.4.8 Summary Parameters for the Safe Design of a Process Cycle General Design Procedures 12.6.1 Front End Prepurification 12.6.2 Reversing Heat Exchangers 12.6.3 Nonreversing Heat Exchangers 12.6.4 High-Pressure Column 12.6.5 Main Condenser Limits of Contaminants and Analysis 12.7.1 Argon as a Contaminant 12.7.2 Propane as a Contaminant Rotating Machines and Other Equipment 12.8.1 Expansion Machines 12.8.2 Liquid Oxygen Recirculating Pumps 12.8.3 Liquid Purge Lines 12.8.4 Liquid Oxygen Disposal Safe Practices 12.9.1 Analytical Equipment Summary Safety in the Design of Dynamic Oxygen Systems Causes of Combustion Test Procedures and Results as Explained by de Jessey The Following Recommendations Are in Order 12.14.1 Flow Velocities 12.14.2 A Very Careful Selection of Materials Consideration of Dynamic Oxygen Conditions 12.15.1 Tests 12.15.2 Example 12.15.3 Nonferrous Metals 12.15.4 Further Studies 12.15.5 Nickel and Its Alloys 12.15.6 Inconel Alloys 12.15.7 Stainless Steels 12.15.8 Copper and Its Alloys 12.15.9 Aluminum Bronze 12.15.10 Aluminum and Its Alloys 12.15.11 Compatibility of Aluminum and Its Alloys for Structured Packings (Structured Packing Consists of Corrugated Strips Coiled and Used as Distillation Trays) 12.15.12 Supplementary Tests on Aluminum 12.15.13 Replication of Aluminum Testing 12.15.14 Machines Used in the Fabrication of Structured (Corrugated) Packing 12.15.15 Summary 12.15.16 Iron Alloys
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12.15.17 Nonmetallic Materials 12.15.18 Lubricants 12.15.19 Caution References For Further Reading Chapter 13 Cleaning for Oxygen Systems 13.1 Overview 13.2 General Considerations 13.3 Cleaning Requirements for Oxygen Systems 13.3.1 Inspection Standards for Fixed Surfaces 13.3.2 Inspection Standards for Movable Parts 13.3.3 Cleaning Procedures 13.4 Equipment other than Piping 13.4.1 Cleaning Procedures: General 13.5 Cleaning Procedures for Carbon Steel Piping 13.5.1 Definition and Recognition of Contaminants 13.6 Cleaning Procedures Available 13.6.1 Blast Cleaning: General 13.6.2 Sand Blasting in Place (Sandjet) 13.6.2.1 Equipment 13.6.2.2 Procedure 13.6.2.3 Inspection and Control 13.6.3 Secondary Cleaning Procedures 13.6.4 Pre-Cleaning before Erection (with Cleaning Reagents) 13.6.4.1 Precautions 13.6.5 Cleaning after Erection: General 13.6.6 Alternative A—with Solvents 13.6.7 Alternative B—Cleaning Agents 13.6.8 Alternative C—with Movable Pistons 13.6.9 Cleaning Stainless Steel and Nonferrous Metals Such as Copper, Associated Fittings, Parts, and Fabrications 13.6.10 Cleaning Aluminum Piping, Fittings, Parts, and Fabrications 13.6.11 Alternate Methods of Cleaning Stainless Steel Pipe, Aluminum Pipe, Copper Tubing, and Their Fittings 13.6.12 Oxygen Compressors 13.7 Cleaning Agents 13.8 Preparation of Cleaning Agents 13.8.1 Caustic or Alkaline Solutions 13.8.2 Acid Solutions 13.8.3 Agents for Stainless Steel, Copper, and Aluminum 13.8.4 Solvents 13.8.5 Aqueous or Semiaqueous Agents 13.9 Drying Gases 13.10 Testing and Inspection Procedures 13.10.1 Indirect Inspection 13.10.2 After Using Solvents or Chemicals
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13.11
Ancillary Tools and Equipment 13.11.1 Wire Brushes 13.11.2 Immersion Tanks 13.11.3 Protective Clothing 13.12 Labor Force 13.12.1 Personnel 13.12.2 Supervision 13.12.3 Inspection 13.12.4 Cleaning Contractors 13.13 Protection and Storage 13.13.1 Piping 13.13.2 Small Equipment 13.13.3 Large Equipment References Chapter 14 Economics 14.1 General Overview 14.2 Historical Background 14.3 Post–World War II Development 14.4 Economic Overview 14.5 Energy Costs 14.5.1 Oxygen Purity 14.5.2 Liquid Oxygen Production 14.5.3 Pure Nitrogen Recovery (Purity at 99.9995%) 14.5.4 Argon and Rare Gas Recovery 14.5.5 Prepurification of Air 14.6 Investment Costs in General 14.6.1 Approximate Allocation of Investment Costs 14.6.2 Contingencies 14.7 Operating Costs 14.8 Maintenance 14.9 Marketing History of Industrial Gases 14.10 Challenging Market Conditions 14.11 Investing in a Project 14.11.1 Raising Investment Capital 14.11.2 Present Value of an Investment 14.11.3 Caveat on the Use of DCFROI 14.12 Envoi Appendix
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1
Gas Separation and Purification of Industrial Gases
1.1 INTRODUCTION Although cryogenic science and engineering fully developed in the twentieth century, it will be fruitful to review and refresh our memory with the various people who dedicated their energies in uncovering the scientific principles and technology, which today form the basic foundations of low-temperature liquefaction and separation of industrial gases, now more elegantly known as cryogenics. One has to go as far back as the eighteenth century to uncover that in 1703 the work of a French physicist Guillaume Amontons in thermometry and mathematics led to the supposition of an absolute zero. Also in 1720, Gabriel Daniel Fahrenheit, born in Danzig, developed the idea of a temperature scale that bears his name, and wherein all the important numbers are divisible by four. In 1741, moreover, Anders Celsius formulated the centigrade scale, better known as Celsius, where zero represents the freezing of water, and 1008 represents the temperature at which water boils. At this point the discovery of dephlogisticated air or oxygen by Joseph Priestly in 1774 should be underlined. He reported his findings to Lavoisier (in France) who wasted no time in repeating Priestley’s experiments and went even further to recognize that it was the active ingredient of the Earth’s atmosphere. Priestley also discovered ammonia, sulfur dioxide, nitrogen, and a gas later identified as carbon monoxide. It was during the nineteenth century that serious attempts were made to liquefy industrial gases and to study their behavior as well as their characteristics. In 1823, Michael Faraday, then a young scientist at the Royal Institution of London, liquefied chlorine, and studied the liquefaction and characteristics of ammonia. He also discovered that ammonia changed from liquid to its gaseous state by absorbing heat from the surroundings, thus lowering their temperature. In 1824, a young French engineer, Sadi Carnot, published his thesis Reflexions, which became one of the pillars of thermodynamics. He demonstrated that a heat engine operating in his proposed ideal, reversible cycle between two heat reservoirs, one hot and the other cold, would be the most efficient engine possible. This ideal cycle is universally known as the Carnot cycle. In 1845, a little known Scot, Thomas Andrews of Queen’s College in Belfast, Ireland, experimented with the change of phase of carbon dioxide at various temperatures and pressures, and generated the first isotherms. These isotherms indicated that a gas cannot be liquefied, regardless of pressure, unless the gas is first cooled to or below a critical temperature. The origin of the absolute temperature scale was due to a British physicist, James Prescott Joule, who experimented with gases in the 1840s to confirm the efficiency of the Carnot cycle. In 1848, he wrote to William Thomson, a colleague, that his deduction was that the efficiency of the Carnot cycle is proportional to 1=T where T is from the equation pv ¼ RT, which represents the combined equations of Boyle and Charles, and T is based on t, the
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actual temperature in degree Celsius plus 273. This sum t þ 273 was then known as the absolute temperature. Then in 1852 Thomson, later known as Lord Kelvin, formalized that the efficiency of the Carnot cycle for an ideal gas is W =Q ¼ (t1 t0 )=(t1 þ 273) ¼ (T1 T0 )=T1
(1:1)
After a long series of experiments with hydrogen, Joule and Thomson announced in 1862 that the difference between the above mentioned absolute scale and the air thermometer was insignificant. After experiments with helium, which was found closer to an ideal gas than hydrogen, 273 became 273.15 with slightly more additional corrections as the temperature dropped below 208C according to Keesom and Tuyen. During 1850, William Thompson (Lord Kelvin) followed by Rudolf Clausius in 1851, articulated the first and second laws of thermodynamics. Around 1865, a young German engineer, Carl Linde, designed a refrigeration machine based on the Carnot cycle and achieved low temperatures by vapor compression. Using the same principle and a spiralwound counterflow heat exchanger (designed by William Hampson of England), Linde built the first industrial air liquefaction machine in 1895. At this point it should be noted that in 1851, an American physician named John Gorrie patented the first ice-making machine using the compression and expansion of air to generate refrigeration1. The patent model is still in the National Museum of American History. In 1873, Johannes Diderick van der Waals developed an equation of state that offered an explanation of the critical phenomena in the change of state, and fitted nicely with the observations of isotherms of Thomas Andrews. Van der Waals used the values of pressure, temperature, and volume divided by their critical values. This led to the law of corresponding states, which enabled James Dewar and Kamerlingh Onnes to determine the necessary data for the liquefaction of gases. In 1877, a Frenchman, Louis-Paul Cailletet2 announced the liquefaction of oxygen. He used a glass capillary tube with thick walls wherein gaseous oxygen at 200–300 bar was cooled to 245 K by evaporating liquid sulfuric acid. After releasing the pressure almost adiabatically, the drop in temperature was sufficient to cause the formation of a mist of liquid oxygen. Almost simultaneously, Raoul-Pierre Pictet3 in Geneva, Switzerland, also liquefied oxygen by compressing oxygen to 200 bar, then cooling it to 173 K by the sublimation of dry ice, which in turn had been cooled by the evaporation of sulfuric acid. In both cases, however, a true continuous liquefaction of oxygen was impossible because the critical temperature of the gas is 155 K. Nevertheless, Pictet pointed the way to a new route to achieve low temperature, namely the cascade principle using a series of refrigerants, each lowering the temperature of the next one. This principle was fully developed by W. Keesom4 at Leyden in his work with helium using ammonia evaporating at 239.8 K, ethylene at 169.5 K, methane at 111.7 K, and nitrogen at 77.6 K. At this point, it is of interest to point out that during the latter part of the nineteenth century three experimental laboratories developed in Europe for the purpose of studying liquefaction, separation, and purification of the industrial gases at low temperatures. The men who worked in these laboratories should be given high credit, because they constantly faced risks from fire and explosion, given the limited availability of safety standards and process equipment at that time. It was only in the latter part of the nineteenth century, for example, that the Royal Institution in England decided to use reverse valve threads to differentiate between inflammable and nonflammable gases. The Royal Institution of London, England, founded in 1799, intensified its low-temperature studies especially under James Dewar, who developed the Dewar’s (vacuum) flask for conserving cryogenic liquids. He liquefied hydrogen as well in 1898, a high achievement at that time.
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In 1882, the University of Leyden, Holland, hired Heike Kamerlingh Onnes4 as professor of experimental physics. During his tenure, Onnes, a close friend of van der Waals, founded the cryogenic laboratory in order to study equations of state, isotherms, and general thermodynamic properties of cryogenic liquids and gases. He became a master experimenter of physics. His laboratory included the development of measuring instruments and the fabrication of material for his experimental work. Postgraduate students worldwide were eager to intern at his laboratory. It was in 1908 that Onnes succeeded in liquefying helium at below 5 K. He also observed superconductivity at 4.19 K. He was unable to solidify helium, however. This latter phase was accomplished by his successor Wilhelmus Keesom in 1926. Keesom then refined the Onnes–Dana data on specific heat, and discovered that at 2.2 K another type of liquid was found, which was named helium II, and it exhibited a smaller density, a greater latent heat of vaporization, and a smaller surface tension—in other words, superfluidity5. Another smaller, but important cryogenic laboratory was founded at the Jagelonian University in Krakow, Poland, around 1876 under the team of Syzgmunt von Wroblewski, a very strong theoretician, and Karol Olszewski3, a capable technologist with a doctorate degree. In 1883, Wroblewski and Olszewski liquefied air, carbon monoxide, nitrogen, and finally oxygen. In fact, with the exception of hydrogen, they liquefied and studied the properties of all the known gases. They also improved the cascade principle that Pictet had pioneered. In January 1884, von Wroblewski produced a continuously changing liquid of hydrogen by cooling the gas with liquid oxygen and allowing it to expand rapidly, a most dangerous experiment by modern standards. In 1885, he also reported a change of conductivity in copper at low temperatures. Unfortunately, in 1914 an invading army closed down the promising laboratory.
1.2 THERMODYNAMICS It has been said many times that ‘‘engineering is the industrial application of science.’’ Unfortunately, many engineers get so involved with technology that they often forget the basic principles of science. In the case of liquefaction of gases and cryogenic systems, however, the two disciplines are integrally associated, and it is very difficult to do any worthwhile work involving cryogenic systems without a comprehensive understanding of the physical branch of science, thermodynamics.
1.2.1 GENERAL PRINCIPLES
OF
THERMODYNAMICS
The First Law or principle states very clearly that the quantity of energy in the universe is constant. It is the law of conservation of energy. Energy cannot be created or destroyed during a process, although it may be changed from one form to another. Indeed, thermodynamics deals with the relationships of various forms of energy. In essence, the First Law states that any change in internal energy Du is independent of the path and may be expressed as uf ui ¼ Q W
(1:2)
where uf is the final energy, ui the initial energy, W the work done, and Q the heat expelled. Energy may be classified into four basic categories: 1. Internal energy (u) due to the motion and configuration of a substance’s molecules, atoms, and subparticles 2. External energy (pv) due to a substance’s pressure and specific volume
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3. Kinetic energy (V2=g) due to the motion of the substance, or system, where V is the velocity and g is the gravitational constant (9.80665 m=s2) 4. Potential energy (Z) due to the substance’s position in space There are other categories such as magnetic and surface energies, but they could be ignored for the time being, as this text deals only with gases. The external manifestations of energy are heat (Q) and work (W). The term heat is applied to the transfer of internal energy from one substance or system to another because of a temperature difference. The term work is applied to the transfer of internal energy to mechanical motion. Neither heat nor work can be considered as internal energy. If a system is subjected to a change, and the final state is identical in all respects to the initial state, and if only work and or heat has been exchanged between the system and its surroundings, then it can be said that . . . .
If the system has received heat, then the system has transferred work If the system has transferred heat, then the system has received work There is a constant relationship between work and heat The empirical relationship between work and heat is
1 cal ¼ 4.1868 J 1 kcal ¼ 4.1868 kJ One calorie is the heat required to raise 1 g of water from 14.58C to 15.58C. One kilocalorie is the heat required to raise 1 kg of water from 14.58C to 15.58C. One joule is the work done by the force of 1 N traveling 1 m in 1 s. It must be emphasized that heat Q and work W are not properties of a system, but rather the effects of the interaction of a system and its surroundings. When the effects of the interaction are balanced, the process is reversible. If they are not balanced, which is more realistic in practice, the process is termed irreversible.
1.2.2
ENTHALPY (H) (J=KG)
Enthalpy, or the total energy of a system, is the sum of its internal and external energies, or (u þ pv) where u is the internal energy and pv is the pressure–volume product. Moreover, the expression (u þ pv) should be considered as an entity and not separately. In the absence of any work other than that of expansion, any addition of heat to the system manifests itself in an increase in enthalpy. Conversely, any extraction of heat results in a decrease in enthalpy. Enthalpy, moreover, is a point property. It has no absolute value; and the engineer deals only with the difference in enthalpy between the initial and the final states of the system. When there is a temperature increase from T1 to T2 at constant pressure, there is an increase in internal and external energies, which can be written as follows: H ¼ H2 H1 ¼ (u2 u1 ) þ (p2 V2 p1 V1 )=J ¼ mcp (T2 T1 )
(1:3)
where m is the molecular weight of the gas and cp is the specific heat at constant pressure. The above equation is applicable to an ideal gas. If the specific heat is a variable and a function of the temperature, the change in enthalpy is an integrated quantity: H¼
ð Final mcp dt Initial
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(1:4)
A change in the function H or (DH) is a measure of all heat effects, and is the basis of all total heat charts and tables. The latter tables are more widely used for the development of cryogenic process systems and operating procedures. To summarize the previous paragraphs, work and heat are manifestations of energy in transition. In a heat exchanger, energy is transferred from a warmer system A into a cooler system B. The energy in transition is heat (Q). On the other hand, if a gas, such as steam expands to drive a piston or turbine, the gas decreases in enthalpy, whereas the machine gains in work. In the latter case, the energy in transition is called work (W).
1.2.3 THE SECOND LAW (THE AVAILABILITY
OF
ENERGY)
The Second Law indicates the limitations of the First Law. The transfer of heat is an irreversible process. It is impossible to transfer heat from one system to another, which has a higher temperature without supplying external work. In general terms, during the transfer of energy, a part of the energy passes through a form that is no longer available for work. But to the engineer a knowledge of energy (enthalpy) is not enough. He must also quantify the availability or irreversibility of that energy for a process. A knowledge of the energy availability or irreversibility in a process, therefore, is mandatory. A short background of the development of the Second Law is as follows.
1.2.4 CARNOT CYCLE The Carnot cycle postulated in 1824 by Sadi Carnot assumes that an engine or a system operating as an ideal reversible cycle between two heat reservoirs (heat sinks) is the most efficient cycle possible. Furthermore, the cycle establishes the upper limit in the efficiencies of all possible cycles or engines. Because there is no change in enthalpy (heat exchange), the net work done by one cycle is equal to the net heat transferred into the system: W ¼ Q1 Q2
(1:5)
Efficiency is W=Q1 or (Q1 Q2)=Q1 or Efficiency ¼ (T1–T2)=T1, where T is in degrees (K). For a machine operating between two different temperatures, one higher at the inlet and the other lower at the exhaust, the maximum quantity of work that can be expected from the machine depends only on the difference between the inlet and exhaust temperatures and is independent of the nature of the fluid. The transferred heat is DQ ¼ (T1 T2 ) where Q is the heat absorbed by the system at T1 and the efficiency of the system is 1 – (T2=T1). From a practical point of view, since T2 cannot equal 0 K, a 100% efficiency is highly improbable for any machine or system. Conversely, a Carnot cycle in reverse, or transferring heat to a system, which has a higher temperature, is the basic process in achieving low temperatures for the liquefaction of gases, and that process demands the supply of external work. At the time, Carnot was completely unaware that the quantity of heat exhausted at T2 was not equal to the quantity of heat supplied at T1. It was about 25 years later, in 1851, that Rudolph Clausius calculated in a purely empirical manner that the quantity of heat exhausted at T2 was less than the heat supplied at T1. He stated, moreover, that this loss of heat (energy)
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was characterized as entropy. Clausius also stated that the concept of entropy was a function of the absolute temperature. The principle of lost work or unavailability of energy can be a most serious challenge to the engineer for the optimization of any given process. According to Dodge6 Clausius’ original statement was: ‘‘It is impossible for a self-acting machine, unaided by an external agency, to convey heat from one body (system) to another which is at a higher temperature.’’
1.2.5
ENTROPY (S)
Entropy is not a property of any system or substance, but rather a mathematical function that quantifies the Second Law. It is a criterion of whether or not a process will take place in accordance with an engineer’s design. Furthermore, entropy is not an energy term by itself. It must be multiplied by the absolute temperature to be equivalent to energy. The product TS is then equal to that portion of energy in a process that is unavailable or unretrievable for further work. In a simpler form expressed by Dodge7: ‘‘It is the least amount of work that will be necessary to restore the system to the state in which it existed before the irreversibility occurred.’’ Dodge also articulated an analogy of entropy as follows: ‘‘If the individual components of the atmospheric air, such as nitrogen, oxygen, argon, etc., existed in individual cylinders, it would be stated that the system was highly organized, therefore, the entropy of that system can be considered as zero, or as a basic point of reference. If, however, the same individual gases were exhausted and mixed together as found in the atmosphere, the system is now highly disorganized, therefore, the entropy of the system has increased considerably, in fact, to the point that the probability of again separating the gases into their original organized state is highly unlikely without external work being applied.’’ Entropy is a useful mathematical function that quantifies the irreversibility of any process. With an understanding of its use, an enterprising engineer can attempt to achieve the maximum reversibility of a system or process. By studying the entropy of every subsystem (i.e., compressor, cooling, distillation, product compressor) of a process and by limiting any probable increase of the entropy of each operation, the engineer can increase the overall efficiency of the process. In other words, the engineer may extract the maximum energy available from the system. Hence the term exergy is now used by many practicing engineers. For example, over the years, engineers working on the low-temperature separation of air, have succeeded in reducing the energy required to produce one metric tonne of contained oxygen in a mixture of 99.5% purity from 350 kW h=h, to about 270 kW h=h and even lower, which represents a saving of 23%. For any reversible process cycle, the algebraic sum of all heat effects divided by the respective absolute temperatures at which the heat transfers occur is equal to zero: Q=T ¼
ð T2
(dQ=T) 0
(1:6)
T1
DS ¼ S2 S1 ¼
ð T2 (dQ=T)
(1:7)
T1
1.2.5.1
Irreversible Systems
Entropy depends only on the state of the system, therefore, the change in entropy depends on the difference of the initial and final states of the system. In an irreversible process, the change in entropy is
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DS ¼
ð Tfinal (dQ=T) > 0
(1:8)
Tinitial
The above expression is an important one, because it indicates that the entropy differential is always positive. Since all separation processes are irreversible, the project engineer, as well as the process engineer have to pay extra special attention to anticipate and prevent any appreciable increase in entropy wherever it may occur, for example (a) Heat Transfer If an amount of heat Q at Th is transferred from a hot reservoir to a cold reservoir at Tc, the latter’s entropy increase is Q=Tc and simultaneously the hot reservoir has a final entropy of Q=Th. The difference in entropy is DS ¼ (Q=Tc ) (Q=Th ) > 0
(1:9)
because Th is larger than Tc, hence Q=Th is the smaller of the two terms. (b) Free Expansion of a Gas If an ideal gas is expanded within an insulated vessel with no work done, and no heat transfer carried out, and with no change in internal energy, then according to the first law Du (internal energy) ¼ Q W
(1:10)
Therefore W¼
ð Vf p dV
and
DS ¼ nRln Vf =Vi
(1:11)
Vi
Since Vf is larger than Vi, the change in entropy is larger. (c) Mixing Process If two gases or liquids for that matter, are mixed wherein the two masses are m1 and m2, the specific heats are c1 and c2, and the temperatures are T1 and T2 with a final temperature at Tf, then the change in entropy will be DS ¼ (m1 c1 ln Tf =T1 ) þ (m2 m2 ln T2 )
(1:12)
One of the terms on the right will always be positive, the other will always be negative, but the positive term will always be greater in value. In summary, therefore, in all reversible systems the change in entropy will always show an increase, and as indicated previously the engineer can only hope to limit the increase to a minimum in each and every item forming part of the overall system.
1.2.6 THIRD LAW Since the liquefaction and separation of gases involves very low temperatures, sometimes in the vicinity of absolute zero, it will be useful to keep in mind the third law of thermodynamics, which was articulated a long time ago by Lord Kelvin: ‘‘At absolute zero, the Carnot cycle cannot expel heat.’’ It was also stated by R. Nernst: ‘‘At absolute zero, entropy is constant.’’ It has been reported that a temperature as low as 0.000001 K has been reached.
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1.2.7
REAL GASES
As stated previously, operations with real gases occur at low temperatures and high enough pressures so that a change of phase (liquefaction) does take place. This variation from ideal gases is due to the presence of molecules, which have volume and are the source of internal energy of the gas. Van der Waals8 succeeded experimentally to demonstrate that the former classic equation of state pV ¼ RT was no longer valid and that a new equation of state had to be developed. He proposed that: (p þ a=V 2 )(v b) ¼ RT
(1:13)
p ¼ RT=(v b) a=V 2
(1:14)
or
where v is the specific volume of the molecules in the gas, and a and b are constants for each gas. Since van der Waals formulated his new equation of state (about 1873), much scientific work has been carried out to modify or make the above equation more precise. Such attempts have proven encouraging, and have also helped the engineer to work efficiently in order to reduce the limitations of the Second Law. It is important to keep in mind that the word volume of a real gas is a measure of the number of molecules of the specific gas that are contained in a given volume or flow. The term normal cubic meter is always measured at 1 atm (101.325 kPa) and at a temperature of 08C (273.15 K). In these conditions, 22.412 normal cubic meters of any gas are equivalent to its molecular weight in kilograms. Since the molecular weight is constant for each component, a mass balance of any system is based on the number of molecules. It can be stated therefore that the mass flow, or weight of any gas depends on the total number of molecules contained in the gas. For example, 1000 g molecules (moles) of air can be separated into approximately 209 moles of oxygen, 782 moles of nitrogen, and 9 moles of argon. With these numbers in mind, it is easy to execute a mass balance of the system. ‘‘What goes in, must come out.’’
1.2.8 1.2.8.1
COMPRESSION
OF
GASES
Critical Temperatures and Pressures
In the nineteenth century, Faraday found that it was impossible to liquefy any gas by compression alone. He opined that the pressure for the liquefaction for any gas must be related to a certain temperature for each gas. Subsequently, Thomas Andrews published the results of his experimental work on the relationship between temperature and pressure for carbon dioxide. His charts demonstrated in graph form the critical temperature and pressure necessary to achieve any change of phase. Much above the critical value, a real gas follows the law for an ideal gas pv ¼ RT; but below the critical temperature, the gas can be liquefied by pressure. Furthermore, the lower the temperature below the critical value, the lower the pressure required for liquefaction (Figure 1.1).
1.2.9
COMPRESSIBILITY
At ordinary temperatures, near 300 K or below, and for small compression ratios, all the known gases with the exception of hydrogen are more compressible than what is indicated
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T1 and higher indicates an ideal gas T2 and lower indicates a real gas T3 is critical temperaturature at critical pressure Pc T4 and lower, gas undergoes liquefaction Pc
C T1 T2 T3 T4 T5
B
A
C is critical point of CO2 To the right of CB fluid is all gas In the area ABC the fluid is mixed phase On the line BC the gas is at its dew point. On the line AC the fluid is all liquid To the left of line AC the liquid is supercooled.
FIGURE 1.1 Isotherms for CO2 at various temperatures developed by Andrews. (From Gomonet, E., Les Basses Temperatures, Production et Emplois, Librairie J.B. Bailliere et Fis, Paris, France, 1952. With permission.)
by Boyle’s law (pV ¼ constant). In other words, the product pV decreases when p increases. Depending on the temperature and the nature of the gas, this decrease reaches a minimum plateau and then begins to increase. At 100 K and below, hydrogen also behaves in a similar manner. (a) Constant pressure change, generally found in cyclic, or continuous flow processes: Q¼
ð T2 cp dT
(1:15)
T1
or p dv ¼
ð V2
p dv ¼ p(v2 v1 )
(1:16)
V1
(b) Isothermal compression (at constant temperature), and for ideal gases, p ¼ RT=V, and the work done is shown in the following equation: ð v2 ð v2 p dv ¼ RT dv=v (1:17) W¼ v1
ß 2006 by Taylor & Francis Group, LLC.
v1
or ¼ RT
ð v2 dp=p v1
or W ¼ RT ln p2 =p1 or W ¼ 2:3026 log p2 =p1 per mole The above equations hold true for compression or for expansion. They also hold true for one only stage, and for one compression ratio, whether the latter is from 1 to 5 bar, or 5 to 25 bar. It is also seen that when an ideal gas is compressed isothermally, the quantity of heat equivalent to the work done must be removed from the gas. The amount of heat to be removed is Q(cal) ¼ W (J)=4:1868
(1:18)
If on the other hand, an ideal gas is expanded isentropically (Joule–Thomson effect across a throttle valve), the heat that is removed is also expressed as Q(cal) ¼ W (J)=4:1868
(1:19)
Dodge has stated that ‘‘isothermal work is the minimum amount of work that must be used to compress a gas over a given pressure range, or conversely, the maximum amount of work to be secured expanding the gas over the same range9.’’ (c) For an adiabatic compression of an ideal gas, the quantity of work required is W ¼ p1 v1 =(k 1)[1 (p2 =p1 )(k1)=k ]
(1:20)
W ¼ RT=(k 1)[1 (p2 =p1 )(k1)=k ]
(1:21)
or
The above equations are for theoretical adiabatic work per mole for either compression or expansion, where p1 is the initial pressure, p2 is the final pressure, and k stands for g the constant for the cp=cv ratio whose value changes for various gases, for example (Table 1.1). (d) In multistage compression, each stage must be treated and calculated as a series of single stage compressors adding the total work for the single machine. As stated previously, the above formulae are entirely theoretical and do not include any mechanical losses or driver inefficiencies.
1.2.10 FREE EXPANSION
THROUGH A
VALVE
At this point it may be profitable to consider the free expansion of a gas through a throttle valve, since its use was the first process to liquefy gases, and is still employed at the present
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TABLE 1.1 Ratio of Specific Heats cp=cv or g for Various Gases Air dry Monatomic gases such as He, Ar Diatomic gases such as O2, N2, H2, CO, Cl2 Polyatomic gases such as CH4, CO2, H2O, SO2 Ammonia (superheated)
1.40 1.67 1.40 1.30 1.31
Source: From Dodge, B.F. in Chemical Engineering Series, McGrawHill, New York, 1944. With permission.
time. It has always been called the Joule–Thompson Expansion named after the early pioneers of cryogenics. Assuming an ideal gas and if only internal and external energies are considered, ignoring potential and kinetic, the effect of a free expansion through a throttle valve may be analyzed as follows: . .
No work is done, so W ¼ 0 No heat is exchanged, so Q ¼ 0
Therefore u 1 þ p1 v 1 ¼ u2 þ p2 v 2 H1 ¼ H2
(1:22)
or DH ¼ 0 As there is no change in enthalpy, an ideal gas cannot be cooled by a Joule–Thompson free expansion. Since there is no change in enthalpy, a Joule–Thompson expansion is called isenthalpic.
1.2.11 INVERSION In a real gas, however, consideration has to be given to van der Waals’ equation of state (p þ a=V2)(v b) ¼ RT, which includes the involvement of the molecules in a gas. The molecules in any gas have finite sizes and their volume is represented in the formula by the letter b. The molecules also display an intermolecular force or attraction, whose value is measured empirically in the formula by the letter a. In the Joule–Thomson effect10 it was found that not all gases showed a cooling effect when expanded through a throttle valve. Certain gases such as hydrogen and helium showed a heating effect. They calculated an inversion point for all gases, which is now known as mJT. The point at which mJT is zero is known as the Joule–Thomson inversion point. If van der Waals’ equation of state is applied to a free expansion of a real gas in order to determine whether or not a gas will be cooled or heated by free expansion, the following equation is useful: mJT ¼ (2a=RT )(1 b=v)2 ={cp [1 (2a=vRT)(1 b=v)2 ]}
ß 2006 by Taylor & Francis Group, LLC.
(1:23a)
The above equation may be simplified for large values of specific volume to mJT ¼ 1=cp (2a=RT b)
(1:23b)
If T is less than 2a=bR, then mJT is positive and the gas will cool. If T is greater than 2a=bR, then mJT is negative and the gas will heat up. In general, those gases whose maximum inversion temperature is greater than ambient, say 290 K, will cool whereas others such as hydrogen, neon, and helium will heat up. In summary, when a real gas undergoes a free expansion through a throttle valve, there is no external work, and no heat is exchanged. It should normally behave as an ideal gas. There is a subtle change, however, in the internal energy of the molecules of the real gas as they disperse downstream of the valve. There is a minute increase in the potential energy of the molecule, which is more than offset by a decrease in its kinetic energy. This phenomenon results in an overall decrease in the temperature, however small, of the expanded gas. 1.2.11.1
Deviation from Boyle’s Law11
Figure 1.2 indicates a diagram for 1 kg of air at various isotherms and pressures. The dotted line reunites the various points of the minimum pv and the dotted diagram somewhat resembles a parabola. Within this dotted area, the air, or for that matter any gas, is subject
11000
0
20
40
60
80
100
120
140
T = 360⬚
7000
280⬚
280⬚
260⬚
260⬚
220⬚
5000
18
0⬚
16
0⬚
4000
200⬚
5000
180⬚
4000
14 0⬚
⬚ 160 ⬚ 140 ⬚ 0 2 1 ⬚ 100 80⬚
12 0⬚
0⬚
10
Va
po
2000
6000
220⬚
⬚
Pv
8000
240⬚
200
3000
9000
7000
240⬚
6000
11000 10000
300⬚
300⬚
8000
200
320⬚
320⬚ 9000
180
⬚ T = 340
340⬚
10000
160
r
3000 2000 1000
1000 id iqu
L
0
0
20
40
60
80
100
120
140
160
180
200
0
FIGURE 1.2 pV=p diagram for 1 kg of air. (From Gomonet, E., Les Basses Temperatures, Production et Emplois, Librairie J.B. Bailliere et Fis, Paris, France, 1952. With permission.)
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TABLE 1.2 Maximum Inversion Temperature for Different Gases Gas
Maximum Inversion Temperture (K)
Helium-4 Hydrogen Neon Nitrogen Air Carbon monoxide Argon Oxygen Methane Carbon dioxide Ammonia
45 205 250 621 603 652 794 761 939 1500 1994
Source: From Barron, R.F. in Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.
to inversion and does not follow Boyle’s law. Along the inversion curve, the Joule–Thomson coefficient is zero because a point on the inversion outline is one at which the slope of the isotherm line is zero. This indicates that outside of the inversion area a gas follows Boyle’s law. For hydrogen or helium, the inversion point is below room temperature. Table 1.2 indicates the maximum inversion temperatures of various gases. It can be concluded from Table 1.2 that most gases can be liquefied by free expansion through a throttle valve, with the exception of helium-4, hydrogen, and neon. The latter gases have to be precooled before being processed with free expansion. Each gas, therefore, requires an individual study of its inversion behavior and the value of its Joule– Thompson coefficient (mJT). For hydrogen, it is necessary to precool the gas to 198 K before any refrigeration can be achieved by the use of a throttle valve expansion. In the case of helium, the precooling will have to reach 45 K.
1.2.12 ADIABATIC EXPANSION Even before Claude successfully applied a reciprocating expansion engine to air separation, the benefits of an adiabatic expansion with external work for refrigeration purposes were already known. It was only a question of how this could be accomplished from a practical point of view. In an adiabatic expansion, the energy in the gas is transformed into external work, and that loss of energy in the gas results in a substantial drop in temperature. It has been stated previously that the final temperature in an adiabatic expansion of an ideal gas is equivalent (in reverse) to the increase in temperature for the adiabatic compression of the same gas (Figure 1.3). The isentropic (adiabatic) expansion of an ideal gas may be stated by the following equation: WAD =m ¼ kRT=(k 1)[(p1 =p2 )(k1)=k 1]
(1:24)
This equation applies to both reciprocating and centrifugal expansion machines. The above expression can also be stated more simply as T2 ¼ T1 (p2 =p1 )(k1)=k
ß 2006 by Taylor & Francis Group, LLC.
(1:25)
Y
T
X
t.
= t. ne co
p
Z
hs
ne
o =c
W
hs = h
con et.
Ri
sin
p
R
is
in
g
g
S
FIGURE 1.3 Isothermal compression, isenthalpic expansion, and isentropic expansion on (T1S) diagram. (From Springman, H. 1985. Cryogenics, principles and applications. Chemical Engineering. (May 13). pp. 58–67.) Note: The curved line from y to z indicates an isothermal expansion by a throttle valve, using the Linde high-pressure cycle. And the line y to w shows an isentropic (adiabatic) expansion with the Claude cycle. (From Springmann, H., Chem. Eng., May 13, 58, 1985. With permission.)
In the above equations, k ¼ g, the ratio of specific heats. In the case of adiabatic expansion of an ideal gas, the principles involved are the same as those for the adiabatic compression, except that the discharge temperature T2 is much lower that the inlet temperature T1. Furthermore, the expression for the work done, W, is negative since the expansion machine is actually the driver. The cooling effect of the adiabatic expansion of an ideal gas is considerable. For example, the expansion of air at 290 K and at 5 bar down to atmospheric can be calculated from the formula: T2 ¼ T1 ( p2 =p1 )0:29
(1:26)
and since T1 ¼ 290 K, and p2=p1 ¼ 1=5, therefore, T2 ¼ 182 K, or 273–182 K. The overall temperature drop is from 290–182 or 108 K. From the above calculations, it is obvious that for an increase in the cp=cv ratio, the lower will be the final temperature in an adiabatic expansion. For example, for a monatomic gas with a compression ratio of 1.67, the final temperature will be 152 K, i.e., a temperature drop from 290–152 or 138 K.
1.2.13 THERMODYNAMIC CHARTS AND TABLES Real gases do not follow any predetermined algebraic equation of state. For this reason, a series of charts and tables have been prepared over the years to represent almost all the thermodynamic properties of real gases in order to facilitate the engineer’s calculations over a wide range of temperatures and pressures. These charts involve the relationships of temperature–entropy, temperature–enthalpy, and enthalpy–entropy. These charts are also called
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Mollier diagrams, which are familiar to most engineers. More recently, however, with the aid of computers many companies in the cryogenic field have developed their own tables indicating even small changes in enthalpy over a wide range of cryogenic temperatures as well as pressures. With the aid of such tables the process engineer can develop a more accurate process design than by dealing with charts and nomograms, which involve a certain amount of guess work. In dealing with charts, and more specifically with tables, it is necessary to keep in mind that there are no absolute values available for the thermophysical properties of real gases. The values of enthalpy or entropy found in the literature have been calculated above a definite reference state. For example in steam tables, the reference state is taken as the saturated liquid at 273.15 K (08C). This can be interpreted as follows: (a) The enthalpy and entropy for this state as both zero or (b) The properties recorded in the steam tables are really changes in the property between the state in question and the reference state Interpretation (b) is correct. In a flow process, where work effects and potential and kinetic energy changes can be ignored, the following equation is valid: Q ¼ hfinal state hinitial state
(1:27)
In cryogenic design, moreover, the reference point, or state may be anything depending on who prepared the chart, table, or software. For example (a) Saturated liquid at normal boiling point (NBT) or (b) Vapor at 273.15 K and 101.325 kPa. (1 bar) For tables relating to air at or near liquefaction temperatures, one must remember that air is a mixture of oxygen, nitrogen, and argon, all of which have different liquefaction temperatures. Some tables involving temperature–enthalpy figures may indicate negative quantities. Nevertheless, the engineer must use them in order to arrive at a proper heat balance in a system where the composition keeps changing. In the calculations of enthalpy and entropy around various points in an air separation system that may involve a composition slightly different than that of air, it will be prudent to calculate the properties of each component separately and to use the total sum as the final property. For example, if the fluid used in the expansion machine comes from the (lower) high-pressure column, it may have a composition different from that of air. When working with charts and tables, the engineer should take care to ascertain the system used for the values stated in the chart. At the present time, there is a certain amount of confusion in the selection of values from either the new International System (SI), or the American system involving ft=lbs. The gas industry standard cubic foot is measured at 608F, whereas the cryogenic industry maintains that a standard cubic foot is measured at 708F. Even in the metric system, a normal cubic meter is always measured at 273.15 K, or 08C, whereas some countries or companies refer to this as a standard cubic meter, and use a temperature of 15.58C, for a normal cubic meter.
1.2.14 CRYOGENIC PROPERTIES OF AIR12 The critical temperature of air is 132.6 K (140.558C) at 37.66 bar (3766 kPa) absolute. Above this temperature air cannot be liquefied by pressure alone. The following general formula may be used for the liquefaction of air.
ß 2006 by Taylor & Francis Group, LLC.
For the beginning of liquefaction log p(bar) ¼ 4:640 0:458(1000=Tk ) þ 0:00733(1000=Tk )2
(1:28)
For the completion of liquefaction log p(bar) ¼ 4:764 0:476(1000=Tk ) þ 0:00733(1000=Tk )2
(1:29)
Between the temperatures of 81.15 K and 78.75 K (192.008C and 194.048C), the liquid produced is a variable mixture of oxygen, nitrogen and argon because each of these elements have a different liquefaction point, and great care must be taken in reading charts and tables.
1.2.15 REFRIGERATION AND LIQUEFACTION SYSTEMS (IDEAL AND REVERSIBLE) It is important to differentiate between refrigeration and liquefaction systems. A refrigeration system or cycle is essentially a closed system operating at a steady rate and where only work and heat are exchanged. In essence, a refrigeration system absorbs heat at a low-temperature level and rejects it at a higher temperature level. The system involves a compressor, a heat exchanger, a throttle (expansion) valve, and a liquid reservoir–evaporator. The system is normally employed to supply a predetermined quantity of refrigeration and at a fixed temperature. The measure of performance of a refrigeration system is called coefficient of performance (COP). The latter term is defined as the energy removed from the source divided by the work required to remove the energy in question. If the energy (heat) to be removed is Q, and work W is required, then the coefficient of performance, or COP is given by the formula: COP ¼ T1 =(T2 T1 ) where T1 is the lower temperature, and in such instances, the COP is greater than unity.
1.2.16 VAPOR COMPRESSION SYSTEMS13 In vapor compression systems using refrigerants such as ammonia, Freons, which are easily condensed when compressed, and are expanded in the dual phase, as in pressure–enthalpy (Mollier chart), which is more convenient to use than a T–s diagram (Figure 1.4). The energy balance is as follows: Heat absorbed Q ¼ HC HB Heat released Q ¼ HD HA Work required W ¼ Q1 Q2 ¼ HC HB HD þ HA ¼ HC HD
(1:30)
COP ¼ (HC HB )=(HD HC )
(1:31)
Since HB ¼ HA
where H is the enthalpy. As a rule, the COP of vapor compression systems is greater than 1. As noted, the simplified expression, COP ¼ T1=(T2T1) cannot be used for refrigerants using cold gases such as nitrogen.
ß 2006 by Taylor & Francis Group, LLC.
Absolute pressure
Liquid region
Two-phase region A
Superheated vapor region E
Condensation
D
Throttle expansion
B
Compression
Evaporation C
Enthalpy
FIGURE 1.4 Mollier or pressure–enthalpy diagram used in vapor compression systems employing condensation by recompression followed by throttle expansion, evaporation. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
1.2.17 LIQUEFACTION SYSTEMS 1.2.17.1
High-Pressure Free Expansion (Isenthalpic Linde–Hampson) System14
In a liquefaction system, Figure 1.3 and Figure 2.3 the refrigeration duty Q is replaced by a continuous withdrawal of the finite liquid m from the reservoir. The uncondensed portion of the gas is returned to the system in countercurrent–heat exchange, and the liquid is removed from the system as product or for further treatment. In a liquefaction system identical to a Linde–Hampson free expansion cycle, the overall energy balance around the simple system of heat exchanger, throttle valve, and reservoir is as follows: mH2 ¼ (m mf )H1 þ mf Hf
(a) (b) (c) (d)
¼ ¼ ¼ ¼
(1:32)
(b) þ (c) enthalpy entering a heat exchanger enthalpy leaving heat exchanger enthalpy leaving with the liquid product
And if the portion liquefied is mf=m, then the liquid yield y is y ¼ (H1 H2 )=(H1 Hf )
(1:33)
Since the terms H1 and Hf are constant, it is obvious that the yield increases as the temperature H2 at the inlet of the primary heat exchanger decreases. That is to say, that by precooling the system, the overall efficiency of the cycle, and the liquid production increases. Incidentally, the same principle is applicable to a refrigeration system. Linde was well aware of this application, and used an ammonia precooling refrigeration cycle in a later development of his high-pressure system to improve the production of liquid.
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1.2.17.2
Claude Isentropic System15
The Claude system (Figure 1.3 to Figure 2.6) replaces the throttle (expansion) valve with an expansion machine, either reciprocating, or centrifugal, which creates work thus eliminating the increase in entropy, at least theoretically. The throttle valve is not completely replaced in order to avoid the formation of liquid within the expansion unit. From Figure 2.6, it can be seen that the gas is compressed isothermally at ambient temperature, then it passes through a heat exchanger to be cooled to a predetermined temperature countercurrent to the uncondensed gas returning from the evaporator–condenser. A portion of the precooled gas is sent to the expansion machine where it reaches its dew point. Then it is returned to the liquefier– exchanger to condense the remaining portion of the process air. The remaining portion of the air is cooled by the returning air from the expander. It is then sent to the throttle valve for free expansion and condensation in the evaporator–condenser, sometimes called the liquefier. The expansion machine is loaded with an oil brake for very small plants, with a motor=generator set for larger plants, or a booster compressor for very large plants. The overall energy balance around the heat exchanger, expansion machine and the liquefier, applying the First Law, is as follows: mH2 ¼ (m mf )H1 þ me (He H3 ) þ mf Hf
(a) ¼ Enthalpy (b) ¼ Enthalpy (c) ¼ Enthalpy (d) ¼ Enthalpy
(1:34)
entering the first exchanger leaving the first exchanger and re-entering the compressor differential across the expander of liquid withdrawn from system
The yield is y ¼ (H1 H2 )=(H1 Hf ) þ r[(H3 He )=(H1 Hf )]
(1:35)
where r is the rate of flow of the refrigerant used in the expander.
1.2.17.3
Precooling Systems16
Auxiliary refrigeration cycles can and have been used with liquefaction systems in order to increase the liquid yield, and to liquefy gases whose inversion temperature is below ambient. Using Figure 2.6a, the overall energy balance is: mH2 þ mr Hd ¼ (m mf )H1 þ mr Ha þ mf Hf
(1:36)
and the total liquid yield y is y ¼ [(H1 H2 )=(H1 Hf ) þ r[(Ha Hf )=(H1 Hf )]
(1:37)
where r is the rate flow ratio of the refrigerant used for precooling. Moreover, since y ¼ mf=m, it can be shown that the maximum yield of liquid is limited by the temperature of the refrigerant.
ß 2006 by Taylor & Francis Group, LLC.
1.2.17.4
Cascade Systems17
The cascade system for the refrigeration or the liquefaction of gases is merely an extension of the precooling system where each refrigeration cycle is precooled by an independent refrigeration cycle. It is the most efficient process in terms of energy requirements, the closest to a reversible process. The reasoning may be explained as follows. The minimum work W required for the liquefaction of any gas is expressed by Wreversible ¼ H T0 S
(1:38)
It is also established that the sum of all entropies leaving a system is always greater than that of all entropies entering the system and the difference between the two sums is a direct measure of the irreversibility of the process. It has also been established that by precooling a liquefaction system, the entropy differential is less as the initial temperature is lowered. In other words, if a gas is expanded in its saturated or liquid phase, it limits its increase in irreversibility considerably (Figure 1.5 to Figure 2.9). The cascade system, therefore, is an extension of the precooled liquefaction cycle wherein only liquids are expanded. The process employs a series of gases with progressively falling boiling points, each of which is liquefied under pressure through the refrigeration effect caused by the evaporation of the one next higher in boiling point. This was first carried out by Kamerlingh Onnes in Leyden, and treated theoretically by Keesom who succeeded Kamerlingh Onnes. In the liquefaction of air, one may begin the refrigeration process with either an ammonia or Freon cycle, then follow with ethylene or ethane, and finally with methane. There is a lower limit to the temperature attainable by this procedure, however, that is set by the fact that the critical temperature of the lowest usable gas must be higher than the triple point of the next higher one. For example, the evaporation of liquid nitrogen, with a
Specific energy consumption (kw h/kJ)
10⫺2
Hydrogen cycle
Nitrogen cycle 10⫺3
CH4 cycle C2H4 cycles NH3 cycles
10⫺4
Minimum energy expenditure (theoretical) Two stages One stage
10⫺5 280
240
200 160 120 Temperature (k)
80
40
FIGURE 1.5 As the temperature decreases, the specific energy consumption in kW h=kJ increases in order to a unit amount of refrigeration (cold). (Courtesy of Linde BOC Process Plants. With permission.)
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triple point of 63.3 K, cannot be used for the liquefaction of neon with a triple point of 24.5 K, or for that matter, hydrogen with a triple point of 14.1 K. The evaporation of liquid nitrogen can and is being used, however, for the liquefaction of neon and hydrogen, but as an independent precooling agent in a separate neon or hydrogen process cycle. Although the cascade process cycle is the optimum selection for the liquefaction of gases in terms of energy savings, it also presents the disadvantages of mechanical complexity and high capital investment as well as operational problems due to leakages from the mechanical connections of the various refrigerants.
1.2.18 SUMMARY 1. The efficiency of any liquefaction system as defined by the ratio of the reversible work of liquefaction to the actual net energy consumed depends on the irreversible entropy increments occurring at various points in the system. 2. The principal sources of such increments are found in the isenthalpic expansions of the gas, and in the wide differences of temperatures at the cold end of the heat exchangers. 3. Precooling a compressed gas before isenthalpic expansion leads to a smaller irreversible entropy change, especially close to the saturation point of the gas. 4. Every effort should be made to expand a fluid in its liquid phase for a minimum change in entropy.
REFERENCES 1. Morse, M.S. 2002. Chilly reception, Smithsonian Magazine, June, pp. 30–32. 2. Gomonet, E. 1952. Les basses tempe´ratures, production et emplois, Librairie J.B. Bailliere et Fils, Paris, France, p. 66. 3. Gomonet, E. 1952. Les basses tempe´ratures, production et emplois, Librairie J.B. Bailliere et Fils, Paris, France, p. 67. 4. Gomonet, E. 1952. Les basses tempe´ratures, production et emplois, Librairie J.B. Bailliere et Fils, Paris, France, p. 172. 5. Timmerhaus, K.D., T. Flynn. 1989. Cryogenic process engineering, International cryogenic monograph series, New York: Plenum Press, pp. 26–31. 6. Dodge, B.F. 1944. Chemical engineering thermodynamics, New York: McGraw-Hill, p. 49. 7. Dodge, B.F. 1944. Chemical engineering thermodynamics, New York: McGraw-Hill, Chapters 2 and 3. 8. Gomonet, E. 1952. Les tre`s basses tempe´ratures, production et emplois, Libraire J.B. Bailliere et Fils, Paris, France, p. 31. 9. Dodge, B.F. 1944. Chemical engineering thermodynamics, New York: McGraw-Hill, pp. 265–267. 10. Baron, R.F. 1985. Cryogenic systems, Second edition, Oxford Science Publications, Monograph on Cryogenics #3, p. 65. 11. Baron, R.F. 1985. Cryogenic systems, Second edition, Monograph on Cryogenics #3, Cambridge: Oxford, p. 69. 12. Gomonet, E. 1952. Les tre`s basses tempe´ratures, production et emplois, Libraire J.B. Bailliere et Fils, Paris, France, pp. 70–71. 13. Dodge, B.F. 1944. Chemical engineering thermodynamics, New York: McGraw-Hill, pp. 430–435. 14. Barron, R.F. 1985. Cryogenic systems, Second edition, Monograph on Cryogenics, #3, Cambridge: Oxford, pp. 69–73. 15. Baron, R.F. 1985. Cryogenic systems, Second edition, Monograph on Cryogenics #3, Cambridge: Oxford, pp. 85–88. 16. Baron, R.F. 1985. Cryogenic systems, Second edition, Monograph on Cryogenics #3, Cambridge: Oxford, pp. 73–77. 17. Baron, R.F. 1985. Cryogenic systems, Second edition, Monograph on Cryogenics #3, Cambridge: Oxford, pp. 83–84.
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2
Industrial Applications
2.1 EARLY DEVELOPMENT OF INDUSTRIAL LIQUEFACTION SYSTEMS1 Early experimentation on liquefaction of gases aroused the imagination of industrialists like Carl Linde who was already in the business of manufacturing refrigeration machines with ammonia as the refrigerant. His company, the Linde Eismachinen AG, was founded in 1871. With his background, Linde foresaw the possibility of recovering large quantities of oxygen from the atmosphere, and using it in the metallurgical industry as an enriching agent to increase operating efficiency in the recovery and production of metals. However, Linde’s initial trials in 1895 were not very successful, however. He tried to liquefy air by compressing it to 800 bar, and then expanding it through a throttle valve. He had hoped to cool the compressed air at a rate of 0.35 of a degree per 1 bar pressure drop. He soon found out that this procedure did not work. (The processed air was not at its critical temperature, and was outside of its inversion curve.) In 1899, having demonstrated the solubility of acetylene in acetone, Georges Claude also became interested in the possible use of oxygen in a low-cost production of calcium carbide.
2.2 HEAT EXCHANGERS In all the early experiments for the liquefaction of air and other gases, a crude apparatus resembling a countercurrent heat exchanger was used to help precool the gas to be processed. The lack of efficiency of this apparatus was a key obstacle to the successful development of a viable cryogenic industry as it is known today. At this point, therefore, it is only fair to give credit to William Hampson2 and to Heylandt for the ingenious design and development of the wound-coil countercurrent heat exchanger (Figure 2.1). This design offered a very efficient method of heat transfer in a highly compact configuration and gave the infant cryogenic industry a real impetus in its passage from the laboratory to a large-scale application in industry. This exchanger involved a large number of small copper tubings spirally wound within a cylindrical shell, also of copper. The incoming warm process air passed through the tubing in countercurrent heat exchange and at right angles to the colder outgoing products that flowed through the shell. With this design it was possible to enclose a large heat exchange surface within a relatively small external shell volume, as well as increase the efficiency of the heat transfer surface. It was the development of this concept that enabled Linde to launch the first industrial air liquefaction machine in 1896. In fact, some people call the process the Linde–Hampson cycle for air liquefaction. Although this heat exchanger design is still being used for many applications, especially for the liquefaction of large quantities of natural gas, it has been supplanted to a large extent by the development in 1950 of the brazed aluminum extended surface plate–fin exchanger, better known as BAHX. The latter exchanger offers an extremely large surface of heat transfer, which can be packed in a relatively small space. It also offers the economic advantage of being less costly to manufacture. This subject is discussed in more detail in a later chapter.
ß 2006 by Taylor & Francis Group, LLC.
FIGURE 2.1 Coiling a subcooler. (ß Air Liquide, all rights reserved, 2006. With permission.)
2.3
EXPANSION MACHINES3
The possible use of an expansion machine for obtaining low temperatures isentropically was first suggested by William Siemens in 1857. In 1885, E. Solvay of Belgium attempted to use a reciprocating expander to operate at very low temperatures, but could not reach any temperature lower than 175 K (988C). The main problem was the lubrication of the machine at such low temperatures. Even Linde thought of using this concept but rejected the idea as impractical at low temperatures. In 1902, it was Georges Claude of France, who succeeded in applying the reciprocating expansion machine to the air liquefaction process. He used a burnt (degreased) leather packing as a piston seal and operated it without lubrication. With this adaptation, Claude was able to achieve an almost isentropic (adiabatic) expansion of the processed air, resulting in a lower temperature than was hitherto possible, with an air pressure of only 40 bar. In 1939, Pyotr Kapitza of Russia improved on the Claude cycle by suggesting the use of a centrifugal turbine as an expansion machine because of its higher operating efficiency. Since then the use of the centrifugal expander has taken over almost 100% of the liquefaction requirements of the cryogenic industry. During the past 20 years, moreover, there has been a tendency to couple the centrifugal expander directly to a blower compressor, and to boost a portion of the prepurified process air (normally 20% to 30%), to develop more energy to offset heat gain and any extraction of liquid product. This additional liquid product may be used either for storage and eventual transport, or for boosting to a higher pressure by means of a cryogenic pump. This eliminates a gaseous product compressor that has a high investment and operating cost.
2.4
CONTEMPORARY LIQUEFACTION CYCLES
The general parameters for the design of liquefaction cycles thus involved compression, prepurification, and precooling to reach the critical temperature for the gas in question, and the expansion through an expansion machine or a throttle (expansion) valve or both.
ß 2006 by Taylor & Francis Group, LLC.
QR
Compressor m
2
1
1
2
We
con st
Expander
f
p=
T
We
f
f m
s (a)
Liquid
Liquid reservoir (b)
FIGURE 2.2 A thermodynamically ideal liquefaction system. (From Barron, R.F., Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.)
The variety of refrigeration process cycles that are available today to the process engineer are too numerous to be listed in detail. Final selection depends on the imagination and ingenuity of the process engineer, and on the constraints imposed by economics. Before outlining the principal cycles in use today, and on which all other cycles are based, it may be useful to review the original Linde, Claude, and cascade cycles in terms of efficiency and economic viability. These basic cycles are still in use along with an infinite variety of modifications and refinements to limit the increase in entropy to a minimum, increase the efficiency, and reduce power consumption and investment. Figure 2.2 shows an ideal cycle using a complete isothermal compression with a complete adiabatic compression.
2.5 LINDE CYCLE (FREE EXPANSION THROUGH A VALVE)4 The Linde cycle was the first successful industrial cycle for the liquefaction of air that involved compression, prepurification, heat exchange, and free expansion through a throttle valve. The air was first compressed at a pressure of 200 bar. In the prepurification phase, the air was first passed through a solution of 15% lump caustic soda in water for the removal of carbon dioxide, and then passed through a tower filled with dry lump caustic for the removal of entrained water and water vapor. As may be imagined, this method of prepurification was not very efficient, and plants had to be stopped almost every week for a complete deriming. Strangely enough, this method of prepurification continued for over 50 years. The purified process air then entered the primary heat exchanger system contained in an insulated ‘‘cold box’’ for the precooling of the process air in countercurrent heat exchange with the outgoing colder products. Downstream of the heat exchanger, the air passed through a throttle valve undergoing a Joule–Thompson free expansion with a slight decrease in temperature. The cooled air was then recycled through the primary heat exchanger system where it cooled down the incoming process air. Obviously, the cooling effect at the start of the process was minimal. After a long period of time, however, the process air coming out of the primary heat exchanger was sufficiently cool and near its dew point (around 100 K), so it became partially liquid after its expansion through a valve. The liquid portion could then be withdrawn, and
ß 2006 by Taylor & Francis Group, LLC.
Makeup gas mf 1 m 1
Q 2
3
1 (m ⫺ mf) g
W
4
f
Liquid
mf
T = const 1
con
p=
p=
con
st
Temperature T
st
2
3
f
h = const
4
g
Entropy s
FIGURE 2.3 A simple Linde–Hampson high pressure and T–s diagram liquefaction system. (From Barron, R.F., Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.)
the uncondensed vapor portion could be recycled to maintain a temperature balance throughout the entire system. In the first Linde cycle, Figure 2.3, the refrigeration available was very small because it operated at almost constant enthalpy (isenthalpic). The small drop in temperature was provided by the kinetic energy of the molecules in the expanding gas, and not from any work done by the gas itself. A thermodynamic energy analysis of the original Linde–Hampson cycle indicates the following.
2.5.1
THEORETICAL ANALYSIS
OF THE
FIRST LINDE HIGH PRESSURE CYCLE5
Assuming a steady flow to the combined unit involving the heat exchanger, expansion throttle valve, and receiver, and applying the First Law, which states that 0 ¼ (m mf )h1 þ mf hf mh2
(2:1)
y ¼ mf=m ¼ (h1 h2 )=(h1 hf )
(2:2)
liquid yield is given by
ß 2006 by Taylor & Francis Group, LLC.
Example In the application of the simple, original, and thermodynamically perfect Linde cycle, and assuming the following gas property values for nitrogen, if the latter is the refrigerant then h1 ¼ 462 J=g at 101.325 kPa and at 300 K h2 ¼ 432 J=g at 20265 kPa and at 300 K hf ¼ 29 J=g at 101.325 kPa and at saturated liquid phase s1 ¼ 4.42 J=g K at 101.325 kPa and at 300 K s2 ¼ 2.74 J=g K at 20265 kPa and at 300 K sf ¼ 0.42 J=g K at 101.325 kPa and at saturated liquid The liquid yield therefore from Equation 2.2 is y ¼ (462 432)=(462 29) ¼ 0:0693 and the work per unit mass compressed from Equation 2.2 is W1=mf ¼ T1 (s1 s2 ) (h1 h2 ) ¼ (300)(4:42 2:74)(462 432) or ¼ 504 30 ¼ 474 J=g
(2:3)
The work per unit mass liquefied is W1=mf ¼ 474=0:0693 ¼ 6840 J=g
2.5.2 THEORETICAL ANALYSIS
OF
LINDE BASIC CYCLE
WITH
PRECOOLING6
In a subsequent modification, Linde made use of a prior development of Joule–Thompson in their formula for expansion of air: Tk ¼ ( p1 p2 )(273=T)2
(2:4)
where the pressures are a few bars, and T is near ambient. If the temperature decreases, the expansion increases the temperature difference. Linde took advantage of this feature by placing an ammonia machine unit between the main air compressor and the primary process heat exchanger. This added unit lowered the temperature of the process air to 2408K (338C). Presently, ammonia chillers have been replaced by other more reliable refrigeration machines that are more environment-friendly. Example As before (see Figure 2.4), the air was compressed to 200 bar, but the primary heat exchanger was split into two sections. Between the two sections, process air entered a refrigeration machine lowering its temperature to 270 K, (338C). The property values from the T–s diagram for nitrogen used as the main refrigerant are the same as for the simple high pressure Linde cycle; ha ¼ 207.94 kJ=kg refrigerant at 101.325 kPa and at 300 K, hb ¼ 250.2 kJ=kg compressed refrigerant at 681.7 kPa and at 372.85 K, and hc ¼ 61.23 kJ=kg refrigerant at 300 K and saturated liquid. In this specific case involving an external refrigeration machine the thermodynamic balance is Yield, y ¼ ½(h1 h2 )=(h1 hf ) þ r½(ha þ hc )=(h1 hf ) ‘‘r’’ is the refrigerant mass-flow rate, or mf=m.
ß 2006 by Taylor & Francis Group, LLC.
(2:5)
Cooling water Refrigerant compressor a mr
b
c
Refrigerant condenser
J−T valve
We2 a
QR
Makeup gas
d
2
4
3
m
1
6
1
J−T valve
Heat exchangers (m ⫺ m f)
We1
g
5
Main compressor
f
Liquid
mf
T = const
2
3
Refrigerant Boiling point
6
4
f
p=
p=
con
st
con
st
Temperature T
1
h = const
5
g Entropy s
FIGURE 2.4 A precooled Linde–Hampson high pressure liquefaction system, and T–s diagram. (From Barron, R.F., Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
The first term is the normal yield for a simple Linde cycle, the second term is the improvement by virtue of precooling. The mass flow rate of the refrigerant is 0.10. Therefore, the yield is y ¼ ½(462 432)=(462 29) þ 0:10½(207:94 61:23)=(462 29) ¼ 0:0693 þ 0:0339 ¼ 0:1032 which indicates an increase of 48.92% by precooling. The work requirement may be calculated from W=m ¼ T1 (s1 s2 ) (h1 h2 ) þ r(hb ha )
(2:6)
Hence W=m ¼ (300)(4:42 2:74) (462 432) þ (0:10)(250:20 207:94) ¼ 504 30 þ 4:2 ¼ 478:2 kJ And the work per unit mass liquefied is therefore W=mf ¼ 478:2=0:1032 ¼ 4634 J=g The improvement ¼ 6840 4634 ¼ 2206 J=g or 33%.
2.5.3 THEORETICAL ANALYSIS
OF THE
LINDE HIGH-PRESSURE DUAL PROCESS7
Linde also took the efficiency of the cycle another step further with the use of a dual compression cycle plus precooling. He employed the simple well-known fact that the compression energy varies as the log10 of the compression ratio. The compression ratio of the original cycle was 200=1 and its log10 was 2.310, whereas in the dual cycle the compression ratio of the main air compressor was 200=50, or 4 and its log10 was 0.6021. The process cycle, Figure 2.5, involved the initial compression of a smaller quantity of the process air, 20%–25% of the total, to a pressure of 50 bar and then recycling this quantity from 50 to a higher pressure of 200 bar. The higher-pressure stream passed through a threechannel heat exchanger and expanded to the basic pressure of 50 bar resulting in the liquefaction of a portion of the stream. The unliquefied portion is returned to the highpressure second compressor after passing through the heat exchanger cooling the incoming stream. Assuming that i is the intermediate-pressure stream flow-rate ratio (i ¼ mi=m) and m is the total mass flow rate through the high-pressure compressor, the yield of the liquid is y ¼ ½(h1 h3 )=(hi hf ) i½(h1 h2 )=(h1 hf )
(2:7)
And the work requirement=unit of mass of gas compressed in the high-pressure compressor is W=m ¼ ½Ti (s1 s3 ) (h1 h3 ) i½T1 (s1 s2 ) (h1 h2 )
(2:8)
Example Using nitrogen as the working fluid between 101.3 kPa (300 K) and 2030 kPa, to determine the liquid yield and work requirement per unit mass. The intermediate pressure is at 507 kPa and its flow rate is 0.80. In these conditions the properties are
ß 2006 by Taylor & Francis Group, LLC.
N2 system Q
W W b Makeup gas
c
Q
m
2
a
mN
2
3
LN2 bath 4
1
7 (m ⫺ mf)
W
5 g
mf
6
Liquid
f
FIGURE 2.5 Linde dual-pressure process cycle. (From Barron, R.F., Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.)
h1 h2 h3 hf s1 s2 s3
¼ ¼ ¼ ¼ ¼ ¼ ¼
462 J=g at 101.3 kPa and 300 K 452 J=g at 507 kPa and 300 K 432 J=g at 2030 kPa and 300 K 29 J=g at 101.3 kPa and saturated liquid 4.42 J=g K at 101.3 kPa and 300 K 3.23 J=g K at 507 kPa and 300 K 2.74 J=g K at 2030 kPa and 300 K
From Equation 2.7 the liquid yield is y ¼ [(462 432)=(462 29)] 0:80 [(462 452)=(462 29)] ¼ 0:0693 0:0185 ¼ 0:0508 The work requirement is W=m ¼ [(300)(4:42 2:74) (462 432)] (0:80)[(300)(4:42 3:23) (462 452)] ¼ 474 277:6 ¼ 196:4 J=g
ß 2006 by Taylor & Francis Group, LLC.
The work per unit of mass is W=mf ¼ 196:4=0:0508 ¼ 3866 J=g As one may see, the double cycle not only indicates a lower liquid yield than the simple highpressure cycle, but also a lower power consumption. Nevertheless, one may add a mechanical refrigerator operating at 240 K (338) and the operating efficiency is increased to 48.7%.
2.6 THEORETICAL ANALYSIS OF THE CLAUDE CYCLE8 The calculation of the Claude cycle may be a little burdensome, but it may prove useful to practicing engineers who may be interested in performing a heat and mass balance around a cryogenic system should an operating problem occur. Following Figure 2.6 and assuming that the expansion engine is reversible and adiabatic, the expansion process is isentropic. In this manner a much lower temperature is achieved than an isenthalpic expansion as practiced by Linde. In the Claude cycle, air or feed gas, is compressed to approximately 40 bar and passed through a primary heat exchanger. About 75% of the feed is diverted, expanded, and returned (reunited) with the cold vaporized gas from the liquid being produced. The combined streams are then recycled and passed through an ultimate expansion valve for total liquefaction. The use of an expansion valve is still necessary to avoid the formation of liquid in the cylinder of the reciprocating expansion machine. In reviewing the T–s diagram and applying the First Law for a steady-state flow through the heat exchangers, expansion valve, and liquid container as a single unit with no heat transfer, then 0 ¼ (m mf )h1 þ mf hf þ me he mh2 me h3
(2:9)
Assuming that x flow passes through the expander, or
x ¼ me=m
the yield is y ¼ mf=m ¼ [(h1 h2 )=(h1 hf )] þ x[(h3 he )=(h1 hf )]
(2:10)
Presuming that the last term is less than unity, the cycle is a definite improvement over the Linde high-pressure process. If the expander work aids in compression, then that work is expressed by We ¼ me (h3 he )
(2:11)
And the net work if the expander is used to aid compression is W=m ¼ [T1 (s1 s2 ) (h1 h2 )] x(h3 he )
(2:12)
It may be concluded that at the point ‘‘3’’ the inlet condition of the expander for temperature and pressure will yield the minimum work required. Example Using nitrogen as an example determine the liquid yield, total work of compression, and the work to liquefy a unit mass of gas for the Claude cycle. The system operates between 101.3 kPa, 300 K, and 5066 kPa. The expander inflow is at 60% and the conditions are at 270 K and 5066 kPa (50 atm).
ß 2006 by Taylor & Francis Group, LLC.
Qr 1
3
2
9 m We
5
4
8
7
me g
We
6
Expander e Evaporator
Liquid
mf
Qa
T = const
2
1
3
p=c
s = const
p=
Temperature T
con
st
ons
t
9
4 8 5
e
7
h= t
ns co
f
6
g
Entropy s
FIGURE 2.6 A Claude low-pressure process cycle using an expansion machine and its T–s diagram. (From Barron, R.F., Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
The T–s diagram for nitrogen gives the following properties: h1 h2 h3 he hf s1 s2 s3
¼ ¼ ¼ ¼ ¼ ¼ ¼ ¼
462 J=g at 101.3 kPa and 300 K 452 J=g at 5066 kPa and 300 K 418 J=g at 5066 kPa and 270 K 238 J=g at 101.3 kPa and se ¼ s3 (Te ¼ 86.1 K) 29 J=g at 101.3 kPa and saturated liquid 4.42 J=g K at 101.3 kPa and 300 K 3.23 J=g K at 5066 kPa and 300 K se ¼ 3.11 J=g K at 5066 kPa and 270 K
From Equation 2.10 the yield is y ¼ [(462 452)=(462 29)] þ 0:60 [(418 238)=(462 29)] ¼ 0:0231 þ 0:2494 ¼ 0:2725 As 0.40% per unit kg was compressed and passed into the liquid vessel, one may assume that 70% was liquefied, and the total work is given by Equation 2.12, or W=m ¼ (300)(4:42 3:23) (462 452) (0:60)(418 238) ¼ 347 108 ¼ 239 J=g compressed The work required to liquefy a unit mass of nitrogen gas is W=mf ¼ 239=0:2725 ¼ 877 J=g liquefied: The above result is very high compared with other systems. Its theoretical operating efficiency has been calculated at 48.3%.
2.6.1 CLAUDE CYCLE WITH PRECOOLING (FIGURE 2.7) The Claude cycle can also be improved by adding a mechanical refrigeration unit immediately after the main air compressor. The calculations are the same as previously noted, but are lowered by dropping the process air temperature to say 278 K. In fact, all designers of present-day air-separation plants use precooling as a standard practice.
2.6.2 CLAUDE CYCLE WITH DUAL PRESSURES7 The basic Claude dual-pressure cycle is similar in principle to the dual-pressure Linde process; but in the Claude cycle, only the portion of the stream entering the expander is compressed to a higher pressure, which reduces the energy consumption considerably (Figure 2.8). Again using nitrogen as the refrigerant, the main process feed is compressed from 101.3 to 1013 kPa (10 bar). Approximately 60% is cooled to 160 K and sent to the expander. The remainder 40% is further compressed to 3530 kPa (35 bar), cooled, and undergoes a J=T expansion through a throttle valve to 101.3 kPa (atmospheric level). Following the T–s schematic, the following properties are derived: h1 ¼ 462 kJ=kg at 101.3 kPa and 300 K h2 ¼ 460 kJ=kg at 101.3 kPa and 300 K
ß 2006 by Taylor & Francis Group, LLC.
Aftercooler Air Compressor
Purification Oxygen Nitrogen
Main exchanger
Expansion engine
Gaseous nitrogen Gaseous oxygen
Expansion valve
Rectification section
Air
Liquid oxygen Liquid nitrogen
FIGURE 2.7 Claude cycle with precooling. (Courtesy of F.G. Kerry, Inc., 1977. With permission.)
h3 ¼ 455 kJ=kg at 3530 kPa and 300 K h11 ¼ 309 kJ=kg at 1013 kPa and 160 K hf ¼ 29 kJ=kg at 101.3 kPa and saturated liquid hg ¼ 229 kJ=kg at101.3 kPa and saturated vapor s1 ¼ 4.421 kJ=kg K at 101.3 kPa and 300 K s2 ¼ 3.732 kJ=kg K at 1013 kPa and 300 K s3 ¼ 3.345 kJ=kg K at 3530 kPa and 300 K s11 ¼ 3.052 kJ=kg K at 1013 kPa and 160 K sf ¼ 0.418 kJ=kg K at 101.3 kPa and saturated liquid Excluding the compressors, the energy balance around the system is me h2 þ (m me )h3 ¼ (m mf )h3 þ me (h11 he ) þ mf hf Assuming that y ¼ mf=m and x ¼ me=m, then
ß 2006 by Taylor & Francis Group, LLC.
(2:13)
Compressor Makeup gas
Heat exchanger QR2
QR1
2 4
m 1
2
3 10
1 11
Wc 1
Wc 2
me
We 9 5 Expander
e
8
6
g
mf
7
Liquid
f
FIGURE 2.8 Claude dual-process cycle. (From Barron, R.F., Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.)
y(h1 hf ) ¼ x(h3 h2 ) þ (h1 h3 ) þ x(h11 he )
(2:14)
The yield is y ¼ [(h1 h3 )=(h1 hf )] x[(h2 h3 )=(h3 hf )] þ x[(h11 he )=(h11 hf )]
(2:15)
The energy balance per unit mass can be computed by adding the energy balance per unit mass of the two compressors and then subtracting that of the expander work, or W=mnet ¼ T1 (s1 s2 ) (h1 h2 ) þ (1 x)T1 (s2 s3 ) (h2 h3 )] x(h11 he ) ¼ T1 (s1 s3 ) (h1 h3 ) x[T1 (s2 s3 ) (h2 h3 )] x(h11 he )
ß 2006 by Taylor & Francis Group, LLC.
(2:16)
From a T–s diagram one finds that at 101.3 kPa, he ¼ 233 kJ=kg and the entropy ¼ 3.052 kJ=kg K. Therefore y ¼ [(462 455)=(462 29)] 0:6[(460 455)=(462 29)] þ 0:6[(309 233)=(462 29)] ¼ 0:1145 The net work is given by W=mnet ¼ 300(4:421 3:345) (462 455) 0:6[300(3:732 3:345) (460 455)] 0:6(309 233) ¼ 203:5 kJ=kg and the ideal value to liquefy nitrogen in these conditions is 768 kJ=kg.
2.6.3
CLAUDE CYCLE
WITH
HIGH PRECOOLING
TO
LIQUEFY HYDROGEN
OR
NEON9
The Claude cycle may be used to liquefy hydrogen or neon without modification by using liquid nitrogen precooling (see Figure 2.9), and assuming the following operating conditions. . .
. . . . .
Para-hydrogen compressed to 0.01 kg=s for liquefier and refrigerator Temperatures of all streams on the warm side of the three-channel heat exchanger are at 293 K Compressor operates between 1.01 and 10.1 bar Inlet temperature to expander is 60.8 K 70% of para-hydrogen diverted through the expander Assume ideality for expander, compressor, and heat exchangers The liquid nitrogen bath operates at a pressure of 1.01 bar (101 kPa)
From existing tables the following properties for para-hydrogen are assumed with reference to Figure 2.9. h1 (101 kPa, 293 K) ¼ 4100 kJ=kg s1 (101 kPa, 293 K) ¼ 64.5 kJ=kg K h2 (1010 kPa, 293 K) ¼ 4100 kJ=kg s2 (1010 kPa, 293 K) ¼ 55 kJ=kg K h4 (1010 kPa, 77.3 K) ¼ 780 kJ=kg h5 (1010 kPa, 60.8 K) ¼ 590 kJ=kg s5 (1010 kPa, 60.8 K) ¼ 82 kJ=kg K h12 (101 kPa, 77.3 K) ¼ 808 kJ=kg hf (101 kPa, 20.4 K) ¼ 256 kJ=kg sf (101 kPa, 20.4 K) ¼ 8 kJ=kg K ha (101 kPa, 77.3 K) ¼ 29 kJ=kg hc (101 kPa, 293 K) ¼ 455 kJ=kg (A) Make an energy balance including the three cold exchangers, expander, J-T valve, and the liquid reservoir to obtain the liquefaction rate of the para-hydrogen:
ß 2006 by Taylor & Francis Group, LLC.
Compressor mN2
Makeup gas
LN2 bath
QR
LN2
m
(m−mf) We me
We Expander
J−T valve
Liquid mf
FIGURE 2.9 Claude precooled process cycle for liquefaction of neon and hydrogen. (From Timmerhaus, K.D., Cryogenic Process Engineering, Springer Science & Business Media, New York, 1989. With permission.)
mh4 ¼ (m mf )h12 þ mf hf þ me (h5 he )
(2:17)
If one assumes that me=m ¼ x, and that mf =m ¼ y, then (h4 h12 ) ¼ y(hf h12 ) þ x(h5 he ) (2:18) therefore, y ¼ [(h12 h4 )=(h12 hf )] þ x[(h5 he )=(h12 hf )]
(2:19)
For an ideal expander s5 ¼ se ¼ 32 kJ=kg K at 101 kPa, and he ¼ 235 kJ=kg and substituting numerical values y ¼ [(808 780)]=[808 (256)] þ 0:7[(590 235)=808 (256)] ¼ 0:260 For 0.01 kg=s compressed, the liquefaction rate ¼ 2.6 103 kg=s.
ß 2006 by Taylor & Francis Group, LLC.
(B) To determine the liquid nitrogen consumption for this liquefaction rate one makes an energy balance around the three-channel heat exchanger and the precoolant liquid nitrogen reservoir. mh2 þ mN2 ha þ (m mf )h12 ¼ mh4 þ mN2 he þ (m mf )h1
(2:20)
m(h2 h4 ) þ (m mf )(h12 h1 ) ¼ mN2 (hc ha )
(2:21)
mN2 ¼ [(h2 h4 ) þ (m mf )(h12 h1 )]=(hc ha )
(2:22)
Substituting numerical values mN2 ¼ [(0:01)(4100 780) þ (0:01 0:0026)(808 4100)]=(455 29) ¼ 2:07 102 kg=s of liquid nitrogen consumed:
2.6.4
THE LOW-TEMPERATURE REFRIGERATOR
The Claude process cycle using nitrogen as a precoolant may also be used as a low-temperature refrigerator by using the vaporization of the liquid produced.
2.7
KAPITZA CYCLE10
In 1939, when Kapitza suggested the use of centrifugal expansion turbines for the Claude cycle, he also suggested the use of reversing regenerators packed with aluminum strips, the latter developed by Linde–Frankl in 1928. These reversing regenerators replaced not only the primary countercurrent heat exchangers but also served as prepurifiers for the process air. They consisted of diagonally corrugated aluminum strips about 5080 mm wide and approximately 3–5 mm thick. These strips were coiled to any diameter and stacked vertically to any height specified by the process engineer. These reversing regenerators replaced not only the primary countercurrent heat exchangers, but also served as prepurifiers for the process air. This combination became the basic process cycle during World War II. Reversing regenerators were widely used for large plants during the war and shortly afterwards, they have since been replaced by brazed aluminum heat exchangers developed in the United States in the late 1940s. Since the war, a wide variety of process cycles based on the basic Claude cycle and the throttle valve have developed. They are too numerous to be described individually but a few specific ones are described in Chapter 3.
2.8
CASCADE CYCLE11
The Cascade process cycle for the liquefaction of gases is certainly not new. For example, in 1890, Kamerlingh-Onnes in Leyden used a series of refrigerants for his experimental work on cryogenics. The refrigerants originally used were as follows: . .
Liquefied methyl chloride at ordinary temperatures between 4 and 5 bar. Liquefied ethylene in a bath of methyl chloride boiling at 183 K and a pressure of 10 mm of mercury. At this temperature ethylene liquefies at a pressure of 5–6 bar. Also ammonia vaporizes at 240 K. Ammonia has been replaced by propane.
ß 2006 by Taylor & Francis Group, LLC.
.
Liquefied ethylene boiling at a pressure of 20 mm of mercury had a temperature of 131 K. At this temperature, oxygen can be liquefied at a pressure of 20 bar. Ethylene has been replaced by ethane.
By vaporizing liquid oxygen at 90 K, it was possible to reach a temperature down to 73 K; and by vaporizing it under a vacuum, one could attain much lower temperatures. Of course, the use of liquid nitrogen is safer and more effective, because it has a lower boiling point of 77.3 K. In an up-to-date cascade cycle, Figure 2.10, the refrigerant gases employed undergo compression, precooling, expansion through a throttle valve, liquefaction, vaporization, and subsequent recycling to their separate compressors. This cycle is an extension of the principles of precooling. The refrigerants are each liquefied and vaporized sequentially in conventional closed loops. Moreover, each refrigerant may be vaporized at several pressure levels in order to achieve lower temperatures, hence increasing efficiency (but at a higher cost of investment and operating complexity).
Cooling water
Ammonia NH3
Ethylene C2H4
Methane CH4
Makeup gas
Liquid nitrogen
FIGURE 2.10 Process Schematic of a classic Cascade cycle using external closed loops of refrigerants: ammonia, ethylene, methane, nitrogen. (From Barron, R.F., Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
In this process, each refrigerant has a progressively lower boiling point than the previous refrigerant used for precooling. For example, nitrogen may be liquefied by using four cycles in series: ammonia boiling at 240 K, ethylene boiling at 189 K, methane boiling at 112 K, and nitrogen boiling at 77 K. This process is thermodynamically efficient because it approximates an ideal reversible cycle. The expansion through a valve occurs at a lower pressure, and therefore results in a very low increase in entropy. In spite of its slightly higher process efficiency, this cycle is seldom if ever used in air separation because of its high refrigerant losses due to leaks, complexity in operation, poorer safety factor, and higher required investment. After World War II, with the tremendous development of the petroleum and petrochemical industries, some engineers did consider these four refrigerants for the separation of air for very large production of oxygen. Calculations confirmed that the cascade cycle was more efficient than the standard Claude cycle. In fact, an engineering company took a patent on it without knowing that it had been conceived and used many years before. In the 1960s, a merchant air-separation plant was built in Louisiana using a cascade cycle of several independent hydrocarbon refrigerants. When the plant went into operation, however, production costs were much higher than estimated due to excessive mechanical maintenance and continuous leaks of the refrigerants. Nevertheless, this cycle has been, and is still being, used for the liquefaction of hydrogen and helium and for the large-scale liquefaction of natural gas. Propane and ethane are abundant in natural gas. Therefore, they can be substituted in place of ammonia and ethylene with a minimal loss in efficiency. The cascade cycle is widely used in the treatment of natural gas for the recovery of pure methane, and is the standard for the large-scale production of liquefied natural gas (LNG). Further discussion can be found in Chapter 8. Chapter 2 has given the reader an overview of the energy required to recover industrial gases. An ideal quantitative summary is given in Table 2.1 at the following conditions of temperature at 300 K and a pressure of 101.325 kPa (1 atm).
TABLE 2.1 The Ideal Work Necessary for Liquefaction of Various Gases Gas Helium-3 Helium-4 Hydrogen, H2 Neon, Ne Nitrogen, N2 Air Carbon monoxide, CO Argon, A Oxygen, O2 Methane, CH4 Ethane, C2H6 Propane, C3H8 Ammonia, NH3
Normal Boiling Point (K)
Ideal Work of Liquefaction (Wl=mf) (kJ=kg)
3.19 4.21 20.27 27.09 77.36 78.8 81.6 87.28 90.18 111.7 184.5 231.1 239.8
8178 6819 12019 1335 768.1 738.9 768.6 478.6 635.6 1091 353.1 140.4 359.1
Source: From Barron, R.F. in Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.
ß 2006 by Taylor & Francis Group, LLC.
REFERENCES 1. Gomonet, E. 1952. Les basses tempe´ratures, production et emplois, 71–91. Paris: Librairie J.B. Bailliere et Fils. 2. Gomonet, E. 1952. Les basses tempe´ratures, production et emplois, 54–79. Paris: Librairie J.B. Bailliere et Fils. 3. Gomonet, E. 1952. Les basses tempe´ratures, production et emplois, 80–87. Paris: Librairie J.B. Bailliere et Fils. 4. Barron, R.F. 1985. Cryogenic systems (Monogram on cryogenics 3), 2nd ed., 61–64. New York: Oxford University Press. 5. Barron, R.F. 1985. Cryogenic systems (Monogram on cryogenics 3), 2nd ed., 69–72. New York: Oxford University Press. 6. Barron, R.F. 1985. Cryogenic systems (Monogram on cryogenics 3), 2nd ed., 73–80. New York: Oxford University Press. 7. Barron, R.F. 1985. Cryogenic systems (Monogram on cryogenics 3), 2nd ed., 91. New York: Oxford University Press. 8. Barron, R.F. 1985. Cryogenic systems (Monogram on cryogenics 3), 2nd ed., 85–89. New York: Oxford University Press. 9. Timmerhaus, K.D. 1989. Cryogenic Process engineering, 132. New York: Plenum Press. 10. Gomonet, E. 1952. Les basses tempe´ratures, production et emplois, 91. Paris: Librairie J.B. Bailliere et Fils. 11. Gomonet, E. 1952. Les basses tempe´ratures, production et emplois, 67. Paris: Librairie J.B. Bailliere et Fils.
FURTHER READING Springman, H. 1985. Cryogenics, principles and applications. Chem Eng (May 13):58–67. Abbot, W.F. 1989. Thirteen ways of looking at the van der Waal equation. Chem Eng Prog (February):25–26. Kenney, W.F. 1989. Current practical appliations of the second law of thermodynamics. Chem Eng Prog (February):57–69.
ß 2006 by Taylor & Francis Group, LLC.
ß 2006 by Taylor & Francis Group, LLC.
3
Air Separation Technology
3.1 AIR SEPARATION OVERVIEW Oxygen is the most prevalent element on the Earth. It constitutes 46.6% by weight of the crust and 89% of the sea water. It also constitutes 23% by weight (20.9% by volume) of the Earth’s atmosphere (see Table 3.1). Oxygen was discovered by Scheele of Sweden (1773), followed by Priestly (1774), and investigated in depth by Lavoisier (1775) who named it oxygen (acid forming in Greek). Lavoisier studied its support of life by inspiration, as well as its acceleration of combustion. Oxygen is a diatomic gaseous element O2, colorless, odorless, and slightly soluble in water (48.9 mL in 1 L at 273.19 K). At 90.18 K and 1 barA, it is a pale blue liquid slightly heavier than water, and at 55.15 K it freezes into a crystalline solid. Both liquid and solid oxygen are paramagnetic, which is rare as other substances are diamagnetic. Though normal oxygen is diatomic, it also exists in the triatomic form as ozone and as monatomic O in the upper atmosphere (Table 3.2). As the industrial revolution accelerated in Europe, especially in the latter part of the nineteenth century, there was an increased demand for primary products such as steel and coal to satisfy production in factories. Engineers and scientists began to review the possibilities of employing oxygen, this newly discovered element that could accelerate combustion, and therefore increase the manufacture of steel products so much in demand. As noted, oxygen was available in the atmosphere but only at approximately 21%. To recover it economically required the design and construction of an industrial-size process unit. Moreover, the only process considered was liquefaction, separation, and revaporization. Gaseous oxygen could increase the smelting of iron ore. Oxygen, in liquid form, when mixed with carbon particles, could be used as a more powerful explosive to mine much needed coal in larger quantities. The liquefaction of air was only the first step to recover and purify large-scale oxygen production. Linde had proved that air could be liquefied on a commercial scale. The next step was to fractionate air and recover oxygen in as pure a state as required by the industry. Linde’s first venture in the separation of air was not very successful. He permitted liquid air to evaporate, assuming that nitrogen (the more volatile component) would evaporate completely leaving behind fairly pure liquid oxygen. Needless to say, the experiment did not work. Even though there was an important difference of 12.9 K in their respective boiling points at 101.325 kPa, a substantial portion of oxygen evaporated along with the nitrogen. In short, Linde found out that separating nitrogen from a mixture with oxygen was not an easy task. Even assuming that atmospheric air is a simple mixture of only two components, namely oxygen and nitrogen, as everyone believed at the time, the separation of these two gases required additional external work. Without wasting any time, Linde therefore turned his attention to distillation and fractionation.
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TABLE 3.1 Constituents of the Earth’s Atmosphere in Mole Fraction Nitrogen Oxygen Argon Carbon dioxide Neon Krypton Xenon Hydrogen deuteride Deuterium Helium
0.78084 0.20946 0.00934 0.00033 1.818 105 1.14 106 8.6 108 3.12 104 1.56 104 1.0 107
Source: Courtesy of Lotepro Data Book, 1975. With permission.
3.1.1
LINDE’S FIRST FRACTIONATION MACHINE
Although Linde had taken out a patent in 1891 for an industrial machine that liquefied air, his first air fractionation column was actually designed only in 1902. It consisted of a single column with random packing. After the usual prepurification to remove CO2 and H2O using NaOH (see Chapter 5), the process air was compressed to 200 bar, and passed through a series of exchangers in countercurrent heat exchange with outflowing products. Air was liquefied using the Joule–Thompson effect. Part of this liquid product was sent to the top of the single tower as reflux. The procedure followed the basic distillation principles known at the time. Unfortunately, the gaseous product leaving the top of the column had an oxygen content of 7%. Calculating the oxygen recovery one finds (Figure 3.1a). 21 ¼ x þ (1 x) 0:07 x ¼ 0:1505 TABLE 3.2 Physical Properties of Oxygen Molecular weight Normal boiling point Vapor pressure Latent heat of vaporization Molar heat capacity of gas @constant pressure in the range of 0–1 bar and at 298 K Thermal conductivity of gas Thermal conductivity of liquid Gaseous real density Liquid density Critical temperature Critical pressure Flammability range Viscosity Dielectric constant
32.00 kg=kmol 90.18 K @ 101.325 kPa 1.52 mbar @ 54.35 K 6.83 MJ=kmol @ boiling point and 101.3 kPa (212.9 kJ=kg) 29.22 kJ=kmol K 0.024 W=m K at 273 K 151.4 mW=m K 1.429 kg=m3 @ 273.15 K and 101.325 kPa 1.141 kg=L at 90.15 K and 1.013 bar 154.576 K 50.4 bar None, but accelerates combustion 188 kg=m s 106 1.4837
Source: Courtesy of Lotepro Data Book, 1975. With permission.
ß 2006 by Taylor & Francis Group, LLC.
N2 + 7% O2
B
Almost pure O2
A (a)
Pure N2
Pure O2 Pump for reduction of pressure
Liquid air
Liquid oxygen
Compressed air (b)
FIGURE 3.1 (a) Linde’s first attempt to separate air. (b) Linde’s second attempt to separate air. continued
ß 2006 by Taylor & Francis Group, LLC.
Liquid 96% N2 and 4% O2
99% N2
C Liquid with about 40% O2
98% O2 D
E
B
Compressed air (c)
A
FIGURE 3.1 (continued) (c) Linde’s final and successful attempt. (Courtesy of Longmans, Green & Co. With permission.)
Recovery of oxygen ¼ 0.1505=21 ¼ 71.75%, which was commercially inadequate. Furthermore, the nitrogen product containing 7% oxygen was unusable. By trial and error, it was determined that a single column required pure liquid nitrogen reflux at the top. This was carried out by sending the liquefied air into the middle of the column. As the liquid entered the column, its vapor rose to the top and passed through a special apparatus containing pure liquid nitrogen boiling at a temperature of 75 K, even though its normal boiling point was 77.4 K. This lower temperature was achieved by boiling the liquid nitrogen at a very low pressure, around 93 mbar. The procedure stripped the rising vapor of all the oxygen, which after being recondensed descended to the bottom of the tower. The vapor leaving at the top was almost pure nitrogen. This process, however, required a vacuum pump with all its ancillary equipment, and also extra apparatus. Eventhough considered an important improvement, the process was not commercially acceptable and was rejected for air separation. But it did open the road for the development of the double column well known today. This development also emphasized the importance of the single column and promoted its refinement for the production of high-purity nitrogen and, ultimately, ultra-high-purity nitrogen so necessary today in the manufacture of electronic components (see Chapter 8). Finally, Claude’s adaptation of the expansion machine to operate at cryogenic temperatures in 1902 completed the general parameters of commercial air separation (Figure 3.1b and Figure 3.1c). At this point, the fundamentals of separation of air merit a review. ß 2006 by Taylor & Francis Group, LLC.
3.1.2 DISTILLATION
AND
FRACTIONATION
Distillation consists of vaporizing a more volatile component from a mixture using heat, and its subsequent recondensation by cooling and separation. The craft dates back to the ancient Alexandrians who produced oil of turpentine by distilling pine resins. In the early 1980s, industry had already developed large-scale stills for recovering alcoholic liquors equipped with bubble cap trays similar to those used today.
3.1.3 FRACTIONATION Fractionation, or fractional distillation, includes the separation of liquids or gases whose boiling points are very close to one another, for example, oxygen and argon. Its technology demands the closest possible contact between the rising vapor of the more volatile component, argon, and the descending liquid of the less volatile component, oxygen, so arranged as to provide the repeated vaporization and condensation needed to enhance separation.
3.1.4 STRIPPING Stripping means the maximum removal of the more volatile component from the liquid by heating, or reboiling, as it may be called. This occurs in the lower section of the upper (low pressure) column, where gaseous nitrogen is stripped from liquid oxygen.
3.1.5 RECTIFICATION Rectification means the enrichment or concentration of the more volatile component (gaseous nitrogen) in the upper section of the (low pressure) column at least in connection with air separation plants, and in the upper section of the (high pressure) column. In the latter case, it relates to the enrichment of liquid nitrogen.
3.2 THEORETICAL CONSIDERATIONS OF FRACTIONATION To separate even a simple binary system in order to recover its components, it is necessary to consider and understand a few basic facts of physics, with regard to vapor pressures and partial pressures.
3.2.1 EVAPORATION
AND
CONDENSATION1
Normally when one thinks of a boiling liquid one has in mind water boiling at 1008C (373.15 K), and its transformation into steam. It can also be easily understood that at elevated altitudes where the atmospheric pressure is lower, water boils at a lower temperature. Alternately, however, in a pressure cooker water boils at a higher temperature. It is a contest between the escaping molecules of the boiling liquid, and the pressure of the molecules of the ambient atmosphere. Figure 3.2a shows the vapor pressure of water as a function of temperature and pressure. If the analogy is carried over to all substances that undergo transformation from solid to liquid to gas and back again, it is clear that at higher temperatures the pressure exerted by the molecules of a vapor leaving a boiling liquid is greater as the temperature increases. The degree of increase in the ‘‘vapor pressure’’ is defined by the vapor pressure as observed in Figure 3.2b. All vapor pressure curves have an upper and a lower limit. The lower limit is known as the ‘‘triple point temperature’’ below which only a solid will exist. The upper limit is known as the ‘‘critical temperature’’ above which liquid cannot exist with vapor. For the latter reason, early pioneers could not liquefy air by compression alone, because they were
ß 2006 by Taylor & Francis Group, LLC.
2.0
0.9
1.8 1.6
0.7
Absolute pressure in. Hg
Absolute pressure (psi)
0.8
0.6 0.5 0.4 0.3 0.2 Vapor pressure of water at 60⬚F is 0.52 in. Hg or 0.26 psia
0.1
20
30
40
50
60 70 80 Temperature (⬚F)
90
50 45 40
1.4
35
1.2
30
1.0
25
0.8
20
0.6
15
0.4 0.2
Absolute pressure (mm Hg)
1.0
10 5
100
FIGURE 3.2 (a) Vapor pressure of water vapor at varying temperatures. continued
unable to cool the air below the critical temperature as required by the laws of physics (see Figure 1.1). Figure 3.3 shows the vapor pressure curves between triple points and critical points of various industrial gases. Figure 3.4 also shows diagrammatically the equilibrium of liquid and vapor concentrations for an oxygen and nitrogen mixture. The lower curve is the liquid temperature curve, also known as the ‘‘bubble point curve.’’ Any one point or temperature in the lower curve is where the first bubble of vapor escapes from the liquid at a given temperature. The upper curve, also known as the ‘‘dew point temperature’’ curve is where the first drop of condensate forms at a given temperature. Figure 3.5 shows the vapor pressure curves of oxygen and nitrogen at various temperatures. It is important to note that at any given temperature, the vapor pressure of nitrogen is always higher than that of oxygen. In other words, nitrogen has a higher volatility than oxygen. It is this difference in vapor pressure between the two gases that provides the basis for the separation of oxygen and nitrogen. The same rule may be applied to all other gas mixtures. Figure 3.4 shows a vapor–liquid equilibrium diagram at a pressure of 1.0325 kPa. If a binary cryogenic mixture is placed in a vessel and allowed to evaporate, the liquid phase is at equilibrium at x, its bubble point at T1 kelvin, and the vapor phase is also at equilibrium at y at T1 kelvin, its dew point. The values of x and y are different and reflect the relative volatility of each component.
ß 2006 by Taylor & Francis Group, LLC.
Critical
1000 800 600
Critical
400 f (r e
200
Pressure (psia)
nt
4) 74
ant
er frig
) 17
t7
id e io x nd o 13 rb nt Ca era g i r f Re
ran
502
Re
e frig
ia
n mo
(re
Am
So lid
100 80 60 40
r
ra ig e
20 10 8 6 4
2 50 nt 22 a t r ge an fri ger 2 e i R fr t1 an r Re ge fri Re
4
11
nt
g
fri
Re
a er
1
t1
an
r ige
fr
Re
13
t1
ge
fri
2
Re
n ra
fri
ge
ra
nt
71
8)
1 0.8 0.6
W
at
er
(re
0.4 0.2 0.1 –100
–80
–60
–40
–20
0 20 40 Temperature (⬚F)
60
80
100
120
FIGURE 3.2 (continued) (b) Vapor-pressure (boiling point) curves of some common refrigerant, to a logarithmic scale. (From King, G.R., Modern Refrigeration Practice, McGraw-Hill, New York, 1971. With permission.)
Figure 3.5 shows a series of phase equilibrium diagrams at various pressures. Note that as the pressure increases, the difference between the upper and lower curves also more or less decreases. It is of interest to note that at a pressure of 5 atm (506.6 kPa), the normal pressure in the highpressure column of an air separation unit, and at 100 K, the content of oxygen in the liquid mixture is approximately 40%, which is about the normal operating practice in the field. At this stage, it is essential to bear in mind several principles that apply to fractionation. As the ancient Alexandrians knew, the concentration of a component in either the liquid or gaseous phase can be made to increase or decrease by simply varying the temperature. The Gibbs phase rule may be interpreted to state that a separation will occur when a temperature gradient is placed across the distillation column. The more volatile components will rise to the top, and the less volatile components will move downward. According to Timmerhaus2 a distillation column is an apparatus that allows two phases to coexist in
ß 2006 by Taylor & Francis Group, LLC.
150 140 130
2
ne
O
tha
110
n
go
Ar
TR
TR
120
Me
q)
(Li
90
TR
100
Air CO 2
–40 –60 1 10−1
10−2 2
2
3
4
3
4
5
1 9 8 7 6 5
9 8 100 7 6
60
TR TR 10−1 2
2
3
4
3
4
5
1 9 8 7 6 5
9 8 7 101 6
100 2
2
3
4
27
9 8
2
1 10 6 9 8 5 7 4 6 5 3
101
(VA)
Air
2
70
TR
80
TR
N
3
80 ⬚F
K
TR
NO
4
60 0
180 160 170
TR
KR
KR KR
3
O
KR
–80
210 200 190
TR
TR TR
TR n to yp Kr
KR
ne
yle
Eth
8 7
–340 –320 –300 –280 –260 –240 –220 –200 –180 –160 –140 –120 –100
230 220
2
CL
TR
KR KR KR KR
e
an
Eth
1 103 6 9 5 8 7 4 6 5 3
–20
250 240
2
SO
XE
Vapor pressure curves between triple points and critical points
1
LF
CC
bar
40
260
a ut Pr
S H2
ne yle op ne r P pa o
CO
2
CL 3 CH
ob Is
2
NH
F
1
L
CC
PSIA
20
300 290 270 280
e en tyl ne Bu uta B
F2 L2 CC
e
ne
e
an
nt
Pe
e Ac
tyl
TR
ne
xa
He
en
T
KR
Be n
KR
ne
ze
ne
ou e
FIGURE 3.3 Vapor pressure curves (between triple points and critical points). (Courtesy of Lotepro Data Book, 1975. With permission.)
equilibrium with one another. Since distillation deals with mixtures of vapor and liquid at the same conditions of temperature and pressure, two other principles are involved: (a) Raoult’s law, which states that p ¼ Po x
(3:1)
where p is the partial pressure of the component in the vapor phase, Po is the vapor pressure of the pure component at the temperature To of the mixture, and x is the mole fraction of the component in the liquid phase.
ß 2006 by Taylor & Francis Group, LLC.
0
Condensation of a binary system Vapor component
100%
Vapor phase
A
y
Dew point
To
D
To
C
x B
Bubble point
Liquid phase
0
Liquid component
100%
FIGURE 3.4 Vapor–liquid equilibrium diagram. (Cryogenic mixture brought to To K and allowed to separate. Liquid phase is at its bubble point and at equilibrium at To K. Vapor phase is at equilibrium at y and at To K. The values of x and y are different and reflect their relative volatility. At the temperature To K, the equilibrium values are x (liquid phase) and y (vapor phase). The ratio of y=x ¼ k (which is the vaporization equilibrium constant). It is this constant at various values of x and y that gives the equilibrium line shown in Figure 3.10. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
(b) Dalton’s law states that the sum of all the partial pressures of all components in a gaseous mixture is equal to the total pressure of the mixture or p ¼ pa þ p b þ p c
(3:2)
where p is the total pressure. Moreover, assuming that the vapor phase in a fractionation column for air separation is an ideal gas, it is possible to equate the ‘‘Y’’ obtained from both Raoult’s Law and Dalton’s Law. Therefore, for all intents and purposes, pY ¼ Po x
(3:3)
where p is the total pressure. In summary, the above equations show that at a given temperature the vapor pressure of nitrogen is always higher than the vapor pressure of oxygen, and that in vapors coming from a boiling mixture of the two components the vapor will always be richer in nitrogen than the liquid mixture. In other words, the nitrogen component is more volatile than the oxygen component. This provides the basis for the separation and fractionation of air.
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135 N2−O2 130 20 atm 125
120 15 atm 115
Temperature (K)
10 atm 110
105 5 atm 100
95 2 atm
90
85
1 atm
80
75
0
0.2
0.4
0.6
0.8
1.0
Mole fraction N2
FIGURE 3.5 Temperature–composition diagram for oxygen–nitrogen mixtures at various pressures. (From Barron, R.F., Cryogenic Systems, 2nd ed., Oxford University Press, New York, 1985. With permission.)
3.2.2
SIMPLE SEPARATION
BY
CONDENSATION
AND
FLASHING (SEPARATORS)
Partial separation of a binary mixture can be obtained either by simple condensation or by evaporation. This can be accomplished quite easily if the two components have a large difference in their degree of volatility (helium–nitrogen), but with more difficulty when the difference in the degree of volatility is close (oxygen–argon). The cooling for separation is generally associated with the expansion through a throttle valve, and the yield in certain conditions is a gaseous and a liquid product with a higher concentration of the more volatile component in the vapor, and a higher concentration of the less volatile component in the
ß 2006 by Taylor & Francis Group, LLC.
Vapor
Feed J−T valve
Liquid
FIGURE 3.6 Separator (flash condenser). (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
liquid product (Figure 3.6). This apparatus is called a separator or a flash condenser, and its throttle valve upstream of the device is adjusted to achieve a specific ratio between the gaseous and the liquid product. Though the design of the device is simple enough, its process calculation requires iteration and patience in the choice of the correct downstream temperature. The degree of separation between two phases can be achieved as soon as the equilibrium conditions exist and may be obtained when yi=xi ¼ Poi=pi ¼ Ki
(3:4)
which is a constant, and is known as the distribution coefficient or more practically, the equilibrium ratio. The above expression shows that the liquid obeys Raoult’s law, and the vapor obeys Dalton’s law, in other words, the relative volatility (ease of separation) between two components is equal to K1=K2. For real (or nonideal) gases, this expression is found through experimentation and is listed in tables as ln Kp=Kpo. In a binary system (oxygen–nitrogen), y1 ¼ K1 x1
(3:5)
y2 ¼ K2 x2 ¼ K2 (1 x1 )
(3:6)
y1 þ y2 ¼ 1 ¼ K1 x1 þ K2 (1 x1 )
(3:7)
and
ß 2006 by Taylor & Francis Group, LLC.
or x1 ¼ (1 K2 )=(K1 K2 )
(3:8)
y1 ¼ K1 x1 ¼ K1 [(1 K2 )=(K1 K2 )]
(3:9)
or
3.2.2.1
Application
For flashing a binary mixture, F ¼ feed, V ¼ vapor, L ¼ liquid, H ¼ enthalpy of vapor, and h ¼ enthalpy of the liquid. The mole fraction of the j fraction in the vapor phase is yj, the mole fraction in the liquid phase is xj, and the equilibrium coefficient for the jth component is Kj ¼ yj=xj. A mass balance of component j around the separator Figure 3.6 will be xFj F ¼ yj V þ xj L ¼ yj (F L) þ xj L
(3:10)
The mole fraction of the jth component in the liquid phase xj ¼ yj=Kj, therefore, Equation 3.10 becomes yj ¼
xF j 1 þ L=F {(1=Kj ) 1}
(3:11)
and the energy (enthalpy) balance is hF F ¼ hL þ HV ¼ hL þ H(F L)
(3:12)
The liquid fraction of the stream leaving the system using Equation 3.12 is L=F ¼ (H hF )=(H h)
(3:13)
Since the sum of the vapor phase mole fractions must equal unity X J
3.2.2.2
yj ¼ 1 ¼
X
xFj =[1 þ L=F {(1=Kj ) 1}]
(3:14)
J
Procedure
To determine the expanded state of a binary mixture .
. .
.
.
Choose a reasonable temperature between the bubble point and the dew point of the mixture in the separator. Calculate y1 from K1=x1. Choose from Table 3.3 and Table 3.4 the K1 and K2 values of the two components at the chosen temperature. Select from Table 3.3 and 3.4 the enthalpies of the feed and of the expanded components at the selected temperature. Calculate the value of L=F from Equation 3.13.
ß 2006 by Taylor & Francis Group, LLC.
TABLE 3.3 Distribution Coefficient Functions, In (Kp =po), for Nitrogen, Oxygen, and Argon (kPa). The Value of the Reference Pressure is po 5 101.3 kPa (1 atm) Nitrogen
Oxygen
Argon
Temperature (K)
101.3
202.6
505.5
101.3
202.6
506.6
101.3
202.6
506.6
78 80 82 84 86 88 90 92 94 96 98 100 102 104 106 108
0.0798 0.3040 0.5282 0.7582 0.9766 1.2008 1.4249 — — — — — — — — —
— — — 0.7048 0.9030 1.1012 1.2994 1.4976 1.6958 1.8939 — — — — — —
— — — — — — — — 1.5503 1.7017 1.8531 2.0045 2.1559 2.3073 2.4588 2.6102
1.3368 1.1164 0.8959 0.6755 0.4550 0.2346 0.0141 — — — — — — — — —
— — — 0.4573 0.3016 0.1459 þ0.0098 0.1655 0.3211 0.4768 — — — — — —
— — — — — — — — 0.6605 0.7877 0.9148 1.0420 1.1692 1.2963 1.4235 1.5506
0.9075 0.7157 0.5238 03319 0.1400 0.0519 þ0.1400 — — — — — — — — —
— — — 0.2516 0.0717 þ0.1082 0.2880 0.4679 0.6477 0.8276 — — — — — —
— — — — — —. — — 0.5518 0.7323 0.8629 1.0434 1.2240 1.4045 1.5851 1.7656
Source: From Timmerhaus, K.D. in Cryogenic Process Engineering, Springer Science & Business Media, New York, 1989. With permission.
TABLE 3.4 Enthalpy of Nitrogen–Oxygen Mixtures. hL 5 Enthalpy of the Saturated Liquid at the Bubble Point; hV 5 Enthalpy of the Saturated Vapor at the Dew Point. 1 J=mol 5 1 kJ=kg mol 101.3 kPa (1 atm) Mole Fraction N2 0.00 0.10 0.20 0.30 0.40 0.50 0.60 0.70 0.80 0.90 1.00
506 kPa (5 atm)
Temperature (K)
Temperature (K)
Liquid
Vapor
hL (J=mol)
hV (J=mol)
Liquid
Vapor
hL (J=mol)
hV (J=mol)
90.2 87.7 85.7 84.1 82.5 81.3 80.4 79.6 78.8 78.1 77.4
90.2 89.5 88.7 87.7 86.7 85.6 84.3 83.1 81.5 79.7 77.4
419 461 519 599 682 779 879 984 1084 1181 1273
7252 7231 7210 7185 7151 7118 7084 7042 6992 6933 6871
108.9 106.3 104.2 102.5 100.9 99.7 98.5 97.4 96.5 95.6 94.2
108.9 107.9 106.7 105.6 104.3 103.0 101.5 100.0 98.4 96.6 94.2
1315 1403 1499 1591 1675 1758 1851 1947 2052 2152 2248
7536 7507 7478 7448 7415 7377 7339 7298 7252 7201 7147
Source: From Timmerhaus, K.D. in Cryogenic Process Engineering, Springer Science & Business Media, New York, 1989. With permission.
ß 2006 by Taylor & Francis Group, LLC.
.
Using summation Equation 3.14 calculate the sum of all components and determine if this adds up to unity. If not, then choose another temperature for the separator. Once the summation equation at the selected temperature adds up to unity or to within the fourth decimal point, find the condition of expanded stream using Equation 3.13 and x2 ¼ (1x1) for liquid components. Use Equation 3.5 and y2 ¼ (1y1) for vapor components.
3.2.2.3
Practical Example (It Requires Iteration)
Assume that a saturated liquid mixture of air of 79% nitrogen and 21% oxygen at 506.6 kPa passes through a throttle (expansion) valve and reaches a pressure of 101.3 kPa. The resulting mixture will be as follows: At 101.3 kPa, the mole fraction of nitrogen will have a dew point of 81.8 K, and a bubble point of 78.8 K. From Table 3.3, the liquid ¼ 932 J=mol and the vapor ¼ 6943 J=mol. The feed stream at 506.6 kPa has hF ¼ 2042 J=mol. Using Equation 3.11 L=F ¼ (6943 2042)=(6943 932) ¼ 0.8153 From Table 3.4 at 80 K and 101.3 kPa K ¼ ln K1 p=po ¼ 0:3040 (nitrogen) K ¼ ln K2 p=po ¼ 1:1164 (oxygen) The distribution coefficients may not be determined: K1 ¼ 1:3553 (nitrogen) and K2 ¼ 0:3275 The right hand side of Equation 3.14 may be determined: 0:790={[1 þ (0:8153)][1=1:3553] 1} þ 0:210={[1 þ (0:8153)][(1=0:3275) 1]}1:0048 þ 0:0785 ¼ 1:0833 > 1 is therefore too large One must repeat the calculations using a lower temperature, say T ¼ 79.07 K With T ¼ 79.07 K, one arrives at h ¼ 1005 J=mol and H ¼ 6916 J=mol; L=F ¼ 0.8309; K1 ¼ 1.2211, and K2 ¼ 0.2956. With these values the result of the right-hand side of the Equation 3.14 is 0:9299 þ 0:0705 ¼ 1:00041 which is near enough to 1. The condition of the exhaust stream is therefore L=F ¼ 0:8309 x1 ¼ (1 K2 )=(K1 K2 ) ¼ (1 0:2956)=(1:2211 0:2956) ¼ 0:761 y1 ¼ K1 x1 ¼ 1:2211 0:761 ¼ 0:929 The mole fraction of oxygen is x2 ¼ 1 x1 ¼ 1 0:761 ¼ 0:239 (liquid) y2 ¼ 1 y1 ¼ 1 0:929 ¼ 0:071 (vapor)
ß 2006 by Taylor & Francis Group, LLC.
3.2.3 FRACTIONATION3 Fractionation of a binary system with two components with close boiling points can be effected by the sequential cascading of condensation and evaporation as explained in flashing, but in one single column. To carry out this concept efficiently, however, it is necessary to ensure that the vapor is in very close contact, and approaches equilibrium with the remaining liquid at every stage of the process. In air separation, the liquid feed is passed in countercurrent with the ascending gaseous mixture. When in equilibrium, the vapor just above the liquid mixture has a higher nitrogen content due to nitrogen’s higher volatility. The vapor passing through the perforations in the trays and the liquid flowing over them tend to remain in equilibrium by exchanging oxygen and nitrogen in contact with each other. This process involves heat exchange, resulting in the condensation of oxygen and evaporation of nitrogen. The ascending gaseous mixture will steadily become richer in nitrogen, and the descending liquid richer in oxygen. From Figure 3.4, it can be seen that a vapor and a liquid in equilibrium always have different compositions. Conversely, if a liquid and a vapor have the same compositions, they cannot be in equilibrium. (In certain azeotropic systems y ¼ x, but this does not apply to air separation.) If one considers the condensation of a binary system where a liquid phase and a vapor phase are brought together in a vessel and mixed thoroughly (Figure 3.4), the two phases will reach a common temperature To. At this temperature the resulting liquid phase will be at its bubble point x, whereas the vapor phase will be at its dew point y. The values of x and y will be different, reflecting the relative volatility of the two components. At temperature To, the equilibrium values are x for the liquid phase and y for the vapor phase. The ratio of y to x is a constant (y=x ¼ k). It is this constant (at relative values of y and x) that results in the slightly parabolic equilibrium line shown in Figure 3.10. Returning to the vapor–liquid equilibrium diagram in Figure 3.4, and drawing a vertical line at any section will show that the vapor at its dew point is always at a higher temperature T1 than T2, the bubble point of the liquid. It can be stated, therefore, that if a liquid and a vapor having the same composition are brought into contact, the compositions of the liquid and the gaseous phase will move apart. The liquid will contain more, and the vapor less of the less-volatile component. This is the basic principle of fractionation. In order to take full advantage of this characteristic behavior of gases at or near their boiling points, it is important to maximize the contact between the vapor and the liquid phases.
3.2.4 FRACTIONATION METHODS A standard fractionation tower consists of a vertical cylindrical vessel that encloses a number of bubble caps or perforated trays or the more recently designed structured packing (see Section 3.3.3 and Section 3.3.4). The liquid feed, containing the components to be separated, is brought in around the middle of the tower and flows horizontally across each tray and in a downward direction until it reaches the bottom sump of the tower. See Figure 3.7a. It is important to maintain the level of the liquid feed at a quasiuniform level across each tray. The vapor that has been stripped from the liquid at the bottom of the tower sump traverses the plates through perforations, and continues upward until it reaches the top of the tower. As the vapor passes up through the thin layer of liquid on each plate, it mixes with the liquid and almost reaches equilibrium before passing upward to the next plate. The plates are so designed as to permit the liquid to pass across and downward from plate to plate. The perforations in the plates are sized to permit a sufficient flow of vapor to be in equilibrium with the quantity of liquid flowing across them. The vapor is supplied with sufficient upward velocity to prevent the liquid from leaking through the perforations.
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Condenser
Heat out
High volatility Low boiling Vapor up
Product out
High volatility Lower boiling constituent
Z
(N2) out
Z9 Y
Mixture in X9
Y9 Feed mixture (air) in
X W
Liquid down
W9
Heat in Boiler (a)
Product out Lower volatility
(O2) out Lower volatility
More volatile component evaporates and joins vapor Liquid layer on plate
Less volatile component condenses and joins liquid (b)
FIGURE 3.7 (a) Typical fractionation column. (Courtesy of Bernstein, J., Cryogenic Consulting Services, Inc., 2006. With permission.) (b) Transformation in the composition of both liquid and vapor during their intermingling on a fractionation plate. As the vapor of the more volatile component traverses the liquid on the plate, heat and mass balance transfer occurs thereby a small portion of the more volatile component in the liquid evaporates and enriches the vapor bubble whereas a small portion of the less volatile component bubble condenses and joins the liquid. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
As bubbles of vapor traverse the film of liquid (Figure 3.7b) heat is exchanged between the two phases. The less volatile component in the vapor condenses into the liquid film on the trays, whereas the more volatile component in the liquid vaporizes and is swept upward in the vapor. Within practical limitations of pressure drop and power consumption, it is important to produce very small bubbles of vapor passing through the liquid to attain optimum conditions of heat transfer and equilibrium balance. If the rate of flow of vapor increases beyond the design point, it will impede the downward flow of the liquid; frothing will occur and the equilibrium balance will be destroyed. If, on the
ß 2006 by Taylor & Francis Group, LLC.
contrary, the vapor flow is below the design point, the liquid will drop through the perforations and dump to the sump. It is important, therefore, to operate the fractionation tower within the design parameters of liquid flow and vapor rate so that a proper mixture of gas and liquid will exist, thereby attaining the correct equilibrium conditions between the two phases. In the sump of the upper (low pressure) tower, the liquid, which is mostly oxygen (the less volatile component), is heated to release nitrogen (the more volatile component). In airseparation plants this heating, generally called reboiling, is accomplished with heat exchange from nitrogen supplied from the lower (high pressure) column. The reboiling oxygen is at a pressure of approximately 1.2 bar and 91.9 K, and the nitrogen is at a pressure of around 5.7 barA and 95.8 K. This difference of 3.9 K is enough to condense the nitrogen and vaporize the liquid oxygen. This condition was normal with the use of vertical copper-tube vaporizer–condensers. Presently, with the use of brazed-aluminum heat exchangers for vaporizer–condensers, a temperature differential of 2 K and even slightly lower has been achieved. The ascending nitrogen vapor comes into intimate contact with a reflux of nitrogen-rich liquid fed down from the top of the tower to complete the enrichment of the vapor. The top section is called the rectification or enrichment section of the tower. The bottom section is called the stripping section. Both terms here apply only to nitrogen, which is the more volatile component.
3.2.5 FRACTIONATION PLATES3 In the past, the number of fractionation plates to be used was usually a trial-and-error proposition. At present, however, designers have several options to calculate the number of theoretical plates to be used for air fractionation, or any other binary system. The options include the analytical method involving plate-to-plate equilibrium calculations, the McCabe– Thiele graphical method, and the Ponchon–Savarit graphical method. The analytical method is obviously the most accurate. Unfortunately, it is also very tedious and requires a lengthy procedure. In addition, there was originally a scarcity of data. Design engineers therefore abandoned the analytical method, and opted for a graphical method, which is simpler and quicker to calculate. The original graphical method was developed by Ponchon and Savarit of France (1922). This procedure was fairly accurate using enthalpy– concentration diagrams. But the lack of proper enthalpy diagrams placed the Ponchon– Savarit method at a disadvantage compared with the McCabe–Thiele method (1925). The latter method uses constant liquid and vapor rates or constant molal overflow throughout the tower. Constant liquid and vapor rates are based on the assumption that the components of the feed have about equal molal heats of vaporization. This assumption is acceptable when the components have comparable molecular structures and similar boiling points. In both methods heat losses are considered negligible. Although the Ponchon–Savarit method may have proved a little more accurate, the McCabe–Thiele method was found easier to use and was almost universally adopted for many applications such as air separation and fractionation. However, it is not accurate enough because enthalpy change or correction is important. Today it has been almost completely replaced by the original plate-to-plate analytical calculations, which, with the aid of fast computers and more precise liquid–vapor equilibrium data, has been rendered quick as well as accurate. It is now possible to calculate in less than 1 min the number of theoretical plates for an air separation column. Moreover, this permits the designer to start with a fixed plate feed or to place the feed at a given plate where the composition is more compatible with that of the feed. Once given the number of theoretical plates it is possible to recalculate very quickly the number of actual plates by varying the efficiency to a more optimum level. Finally, the number of plates may be altered with input from experience with similar process applications. Nonetheless, it may be of interest to have an overview of the older method in order to visualize the parameters involved.
ß 2006 by Taylor & Francis Group, LLC.
Prior to a discussion of the number of plates required for the fractionation of air some definitions are in order. First of all, a fractionation tower is rated according to the number of theoretical plates contained, as well as according to plate efficiency. A theoretical plate may be defined as one that produces an improvement in composition as existing at equilibrium between a liquid mixture leaving a tray and the vapor rising from the same tray. Plate efficiency is defined as the ratio in percent of the number of theoretical plates to the actual number of plates required for fractionation. Since ideal conditions for equilibrium cannot be achieved in practice, one has to assume a certain plate efficiency. If the vapor has an initial composition of yo, and has a composition of y1 as it leaves the plate, compared with an ideal composition of y2, then the plate efficiency is Eff ¼ (y1 yo )=(y2 yo )
(3:15)
In practice the design engineer uses a specific efficiency for a given tray design. This efficiency may vary from one type of tray to another and from system to system, but an overall tray efficiency of 60%–65% or better is not uncommon for air separation systems. Finally, there is an experimentally proven relationship between the feed state (whether it is a subcooled liquid and a saturated liquid, a mixture of saturated vapor and saturated liquid, a mixture of saturated vapor and superheated vapor) and its composition, the two product compositions, the quantity and quality of reflux, and the number of theoretical plates. Any one of them can be calculated if the others are known.
3.2.6
ANALYSIS
OF
FLOW
IN
EQUIPMENT
Referring to Figure 3.8, Section A will apply to the upper or fractionation section of the tower, and B will apply to the lower or stripping section. V represents vapor, L the liquid, D the upper vapor product, B the bottom liquid product, and x and y are components. 3.2.6.1
Case I: Analysis of a Low Pressure Column
Section A (Rectification) Overall Material Balance Total balance is written as V ¼LþD
(3:16)
Vy ¼ Lx þ Dxd
(3:17)
y ¼ (L=V )x þ (D=V )xd
(3:18)
Component balance is written as
and Equation 3.17 can be written as
Equation 3.18 is the equation of the upper operating line. It defines the relationship of y and x between plates.
ß 2006 by Taylor & Francis Group, LLC.
Case I: Fractionation (above feed plate) Case II: Stripping (below feed plate)
C
D
F —Feed mol/h D —Vapor product mol/h B —Liquid product mol/h
A section rectification
C —Condensing system y = Vapor composition more volatile component V
x = Less volatile component
L
liquid composition
V = Vapor L = Liquid
F V
L
V
L
Feed
B section stripping
Heat source reboiling
B
FIGURE 3.8 Analysis of low-pressure column. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
Section B (Stripping section) Overall Material Balance where F ¼ Liquid Feed, L ¼ Liquid descending from Section A. Feed is in liquid phase. Therefore, liquid feed below feed plate is given by LþF
(3:19)
Vy ¼ (L þ F )x (Bxb )=V
(3:20)
y ¼ [(L þ F )=V ]x (B=V )xb
(3:21)
and component balance is
or
ß 2006 by Taylor & Francis Group, LLC.
Equation 3.21 is the equation of the lower operating line. To determine the point of intersection, y has the same value for Equation 3.18 and Equation 3.21 (equating the laws of Raoult and Dalton). Therefore, (L=V )xi þ (D=V )xd ¼ [(L þ F )=V ]xi (B=V )xb
(3:22)
where xi is the value of x at the intersection. Solution of Equation 3.22 gives the value of xi at the intersection when y is the same. (L=V )xi þ (D=V )xd ¼ (L=V )xi þ (F=V )xi (3xb )=V
(3:23)
(D=V )xd ¼ (F=V )xi (Bxb )=V
(3:24)
xi ¼ (Dxd þ Bxb )=F
(3:25)
Fxf ¼ Dxd þ Bxb
(3:26)
or
or
Overall balance states that
Substituting the value of F from Equation 3.25 into Equation 3.24 xi ¼ xf
(3:27)
Hence the operating lines intersect at x ¼ xf. 3.2.6.2
Case II: Consideration of Vapor Feed
As in Case I, the upper operating line is Vy ¼ Lx þ Dxd
(3:28)
But in Case II, the vapor line is below the feed plate and the operating line is below the feed plate. So (V F )y ¼ Lx Bxb
(3:29)
For the point of intersection what is the value of y when x has the same value in Equation 3.28 and Equation 3.29? Let yi ¼ y at intersection so that xi ¼ (Vyi Dxd )=L ¼ [(V F )yi þ Bxb ]=L
(3:30)
Vyi Dxd ¼ Vyi Fyi þ Bxb
(3:31)
or
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Hence Fyi ¼ Dxd þ Bxb
(3:32)
yi ¼ (Dxd þ Bxb )=F
(3:33)
Fyf ¼ Dxd þ Bxb
(3:34)
or
But overall balance is given by
Substitute value of F from Equation 3.34 into Equation 3.33 yi ¼ yf
(3:35)
So the operating lines intersect at yf. Note: The slope of the lower operating line is larger for a vapor feed than for a liquid feed when the same condensation duty is used (see Figure 3.9). The height of a packed column divided by the number of theoretical plates is referred to as the height equivalent to a theoretical plate, or simply the HETP of a distillation tower.
MC
D Nitrogen 99.99% MC = Main condenser B = Rich liquid mol/h D = Liquid nitrogen mol/h F = Feed mol/h V = Vapor V
L
L = Liquid
F (Feed) yf B 40% Oxygen
FIGURE 3.9 Analysis of high-pressure column, Case III. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
3.2.6.3
Case III: Analysis of the High-Pressure Column
The case for the high-pressure (lower) column (see Figure 3.9) may appear a little easier than for the upper column. It should be remembered that the feed may not always be in vapor form: at times the feed into the bottom of the tower may be part vapor and part liquid depending on the process cycle involved. In this case, however, the feed is considered to be all liquid. Therefore, the component balance is Fy ¼ Dx þ Bxb
(3:36)
Fyf ¼ Dxd þ Bx
(3:37)
which is the inter-plate balance.
which is the overall balance. The bottom point of the operating line is given by y ¼ yf and x ¼ xb, since the value of yf from Equation 3.36 and Equation 3.37 gives xb and yf as points on the general operating line. Operating pressure in the tower also plays an important role in the calculation of the number of trays. A lower operating pressure increases the relative volatility and it becomes easier to separate the components, which means fewer theoretical plates or, for the same number, a higher purity for the more volatile component. To determine graphically the number of theoretical plates required, it is first necessary to find out what quantity of nitrogen to be retained in the liquid oxygen sump of the tower as specified by the designer. It is also required to know the purity of the nitrogen to be removed as product at the top of the distillation tower. This data is applied on Figure 3.10. The mole fraction of nitrogen in the liquid phase in the sump of the tower is on the x-axis. The mole fraction of nitrogen in vapor phase at the top of the tower is on the y-axis. A straight diagonal line is the theoretical operating line. The theoretical operating line represents operation of the tower at total reflux, or zero tower output (or feed). The curve joining the operating line at its ends is the equilibrium line, and is somewhat parabolic. At the left of the equilibrium line is the composition of the saturated vapor leaving the (nþ1)th plate, and to the right of the equilibrium line is the composition of the liquid at its bubble point also leaving the (nþ1)th plate at the same temperature as that of the vapor on the other side of the equilibrium line. The slopes of the operating lines are given by Equation 3.18 and Equation 3.19 as are also the intersections of the operating lines with the diagonal. Once the operating lines are drawn, the theoretical plates can be included in the diagram. Beginning at the value of x ¼ 0.95 (95% of vapor nitrogen in the product at the top), a vertical line is drawn from the equilibrium line to the upper operating line. This is the first plate. From this point, a horizontal line is drawn from the operating line to the equilibrium line, and a step formation follows until the steps are past the intersection of the two operating lines, the upper and the lower. With regard to the lower stripping section the same procedure is followed, but beginning at the x value of 0.5, which is the nitrogen content in liquid phase in the sump of the column, proceeding upward. At the end it is found that the stripping section requires five plates and the rectification section requires about 10 plates. It is also of interest to note that as the purity of the components gets higher, the number of graphical steps becomes greater, with each step producing smaller increments in rectification.
3.3
PRACTICAL CONSIDERATIONS
Obviously, the more the number of plates used in fractionation, the higher the purity of the vapor product. For example in air separation for very high-purity nitrogen, it is not unusual
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B
Condenser
on
0.8
ti fica
line
cti
Re
Mole fraction of N2 in vapor
I Feed line C
lin
e
Eq
uil
ibr ium
lin
e
0.6
e
pin
g
0.4
g
rip St
ra
pe
r eo
Th
0.2
0
o al
ic
et
A
lin
tin
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9 0.95%
Mole fraction of N2 in liquid Air separation product nitrogen 95% Product oxygen = 95%, N2 = 5% Stripping section requires 5 theoretical plates Rectification section requires 10 theoretical plates
FIGURE 3.10 McCabe–Thiele graphical analysis for theoretical plates. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
to install close to 90 plates in the upper column. The use of a high number of plates also carries with it the disadvantage of a higher pressure drop, which however means a higher power consumption.
3.3.1 BUBBLE-CAP TRAYS The original bubble-cap tray used many years ago had certain advantages. It was cheap to construct, offered a very good liquid phase level, and had a good liquid retention whenever there was a wide variation or stoppage in vapor flow. Unfortunately it had low efficiency, vapor diffusion in the liquid was poor, and the cost became prohibitive with the increase in air-plant capacity (Figure 3.11).
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Liquid flow
Vapor flow
Downcomer
Bubble-cap plate
FIGURE 3.11 Rectification column plates (or trays). (Courtesy of Bernstein, J., Cryogenic Consulting Services, Inc., 2006. With permission.)
3.3.2
SIEVE TRAYS
The bubble cap soon gave way to the more efficient perforated sieve tray fabricated with aluminum. It was more adaptable as well as more economical for high capacity plants—over 300 t=d of oxygen. The variety of designs for sieve trays are too numerous to describe. They have been used for oxygen plants with a capacity of over 2500 t=d. Even these trays, however, were rendered inefficient and costly for oxygen plant capacities over 1000 t=d. With a tray diameter of 4–5 m it became difficult and costly to construct a sieve tray that could hold a liquid phase absolutely level. In other words it was difficult to maintain a good equilibrium between the vapor and the liquid phase. The individual cost of such trays was anywhere from $2000 to $3000 per tray, and often higher (see Figure 3.12). Sieve trays with a diameter of up to 5 m have been installed in very large air separation plants.
3.3.3
STRUCTURED PACKINGS
In the recent past, however, designers have turned to the use of structured packing for large air-separation plants especially when high power cost was involved. The technology of packed towers is not new. (The use of random packing consisting of ceramic or plastic pall rings has been around for decades.) Structured packing has been used in refineries and chemical plants for a good number of years. It is only since 1980, however, that the industrial gas industry Liquid flow
Vapor flow
Downcomer
FIGURE 3.12 Perforated plate. (Courtesy of Bernstein, J., Cryogenic Consulting Services, Inc., 2006. With permission.)
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has ventured to use structured packing for both lower (high pressure) and upper (low pressure) distillation columns in air-separation plants. The main objective has been to lower pressure drop across the columns to reduce overall energy consumption. At the same time, a well-designed, structured packing will also improve the mixing of the liquid and vapor phases for better rectification. The results from the field since 1990 have been encouraging. Pressure drops have been lowered dramatically and rectification has improved impressively. The technology of structured packing is quite simple and fairly well known. Thin strips of aluminum or copper of a predetermined width are perforated, and then corrugated. The strips may be assembled either in a parallel manner, or coiled in a circular configuration. In either case, their overall geometry resembles a tray. Finally, these trays are installed in the columns either in the shop or at the field. Designers have not yet agreed as to which type of assembly is better. The liquid reflux is fed from the top of the tower and distributed uniformly to the trays below by specially designed distributors. As packed towers are unusually tall, the designer has to consider the use of a number of distributors, as many as five in the lowpressure column to remix the liquid and to correct any tendency to develop poor uniformity of distribution. The functioning of the structured packing tray is similar to that of a reflux condenser, also known as a ‘‘dephlegmator’’ designed by Claude in the 1920s. The reflux condenser is still in use today especially for the recovery of crude argon. Refer to Figure 3.17 in Section 3.7.6. As a film of liquid reflux descends the vertical surface of the tray, it comes into intimate contact with the rising vapor and reaches equilibrium with it. As with a standard sieve tray, heat exchange occurs between the two phases. The more volatile component (nitrogen) in the liquid vaporizes and joins the ascending vapor, whereas the less volatile component (oxygen), entrained by the rising vapor, condenses and joins the descending liquid toward the sump. This structured packing design permits the production of a very thin film of descending liquid in intimate contact at all times with a continuous flow of minute vapor bubbles throughout the entire packing, with a minimal pressure drop. In fact, assuming an ideal design, if one were to take a vertical slice across the tray, one will find a fairly constant thickness of the film and a constant vapor-flow throughout the entire volume, which is the ultimate objective of the rectification process. The height of a packed tower divided by the number of theoretical plates is referred to as the height equivalent to a theoretical plate, or simply as the HETP of a distillation tower.
3.3.4 CARE
IN THE
DESIGN
OF
STRUCTURED PACKING
The use of structured packing, however, is not without certain disadvantages. The design imposes certain difficult demands: surfaces of the packing must be completely wet at all times, with a uniform liquid distribution throughout the entire ‘‘tray.’’ The packing has to be tightly fitted against the side wall of the tower to prevent leakage, and distributors have to be designed accurately and located at the right elevations in the tower. Moreover, it has to be kept in mind that the overall density of the packing per volume of tower is roughly 3.75 times that of an equivalent standard tray, which results in a heavier as well as a taller tower. In short, the use of structured packing, at least for the present, may be limited to air separation plants of a large capacity, 300 t=d and over, and in situations where power costs are much above average.
3.3.5 SAFETY IN
THE
USE OF STRUCTURED PACKING
Though it has been reported to the industry that the use of structured packing made from aluminum may lead to combustion, especially in the lower section of the upper column of an air separation plant; this has been examined in detail by the industry and has been found untrue. Details on safety can be found in Chapter 12.
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OPERATIONAL CONTROL4
3.4
Normally, there are three operating control valves (see Figure 3.13) that are used in the operation of a standard air separation plant.
3.4.1
FCV-1
FCV-1 controls the flow of the liquid nitrogen reflux that is sent from the (lower) high-pressure column to the upper tray of the (upper) low-pressure column. By throttling the valve, the flow is reduced and the content of oxygen is also reduced. This action allows more liquid nitrogen reflux to remain in the high-pressure column reducing entrainment of oxygen. However, if the throttling is overdone, reducing the quantity of reflux available to the upper tray in the upper low-pressure column, it reduces the recovery of the oxygen product. If the valve is opened more than enough, the quality of reflux available to the upper column is excessive, and recovery is reduced because oxygen is entrained by the increase of waste nitrogen. Reflux 450 N2 @ 1% O2 FCV-1 Controls flow of reflux from HP to LP Maintains material balance
800 Waste N2
200 Pure O2
Lower LP column
Reboiler
FCV-3 Controls flow O2 Fine tunes balance
Upper LP column the condenser of the HP column also acts as the reboiler of the column
Condenser
1000 550 FCV-2 Maintains level in HP sump Prevents vapor transfer from HP to LP Lower temperature of RL
1000 mol air
Flow = 1000 – 450 = 550 mol Rich liquid @ 37.36% O2
O2 in RL = 0.21 x 1000 − 0.01 x 450 1000 – 450 = 37.36%
FIGURE 3.13 Diagrammatic view of the distillation of air (operational control). (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
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3.4.2 FCV-2 FCV-2 maintains a constant level in the high-pressure column by throttling the rich liquid from the high-pressure column down to the operating pressure of the low-pressure column. This valve may also be of the ‘‘smart’’ type, with a programmed chip insert. This valve once set, seldom requires any further resetting.
3.4.3 FCV-3 FCV-3 controls the flow of the gaseous oxygen product leaving the upper column. This valve has to be monitored carefully so as not to prejudice the designed operation of the lowpressure column. An excess flow will lower the purity of the oxygen product. A decrease in the flow will increase purity but will also lower recovery of the oxygen product.
3.4.4 REFRIGERATION The quantity of refrigeration must also be monitored closely. Any excess cooling supplied to the plant will manifest itself in increasing the liquid level in the main condenser. It will then be necessary to take out some oxygen product as liquid, and reduce the flow of gaseous oxygen. Alternatively, one might extract part of the liquid nitrogen reflux out of the plant as an additional product.
3.5 PRODUCT RECOVERY The extraction efficiency of an air separation plant is based on its main product recovery. In terms of oxygen product, it may be described as h, the ratio of usable oxygen produced divided by the oxygen in the process air treated by the plant. The above definition can be verified very quickly by checking the quantity of oxygen entrained by the waste nitrogen leaving the low-pressure (upper) column. This observation will quickly verify if the plant is operating within the designed optimum parameters. In order to increase the oxygen-product recovery, one must increase the reflux at the top of the upper column, but this depends on the following: (a) The quantity and quality of liquid nitrogen reflux sent from the lower (high pressure) column (b) The quantity and quality of the vapor rising to the top tray of the upper (low pressure) column (c) The vapor rising to the top tray must always be close to equilibrium with the liquid at the top tray As can be seen from the examples shown diagrammatically in Figure 3.13, the recovery of oxygen product can be increased by throttling the reflux valve slightly and increasing the quality of the reflux liquid, but this procedure has a limit. Although it improves the quality of the reflux by lowering the oxygen content, it also lessens the quantity of reflux. If it is overdone it may decrease recovery. The operator has to be very careful in fine-tuning the recovery (Figure 3.14).
3.6 OPTIMUM REFLUX In reviewing a material balance of the lower (high pressure) column of an air-separation plant, it will be noted that the rich liquid accumulating at the bottom is about 550 mol per 1000 mol of the air feed. Since the total quantity of oxygen is to be recovered, the oxygen
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Case I: Basic operative case 800 KX%
[email protected]% O2
[email protected]% O2
[email protected]% O2
In 400 • 0.015 800 • 0.019
Out 400 • X 800 • KX
6 + 15.2 = 800 X X = 0.0265 % O2 in WM2 = 0.0265 K = 0.01325 = 1.325% O2
Case II: Reduce reflux to improve quality and recovery of product Result: Recovery of O2 has improved 800 KX %
800@2% O2
350@1% O2
350@X% O2
In 350 • 0.01 800 • 0.02
Out 350 • X 800 • KX, K = 0.5
3.5 + 16 = 750 X X = 0.0260 % O2 in WM2 = 2.60 K, K = 0.5 = 1.30% O2
Case III: Reduce reflux still to improve quality Result: Recovery of O2 has decreased 800 KX %
300@5% O2
[email protected]% O2 300@X% O2
In 300 • 0.005 800 • 0.025
Out 300 • X 800 • KX, K = 0.5
1.5 + 20 = 700 X X = 3.07 % O2 in WM2 = K(3.07) K = 0.5 = 1.54%
FIGURE 3.14 Relationship of quantity and quality of reflux on product oxygen recovery. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
content is about 210 mol or 38.2%, the balance being 8.9 mol or about 1.62% of argon, and 331.1 mol or 60.2% of nitrogen. This quantity of 550 mol is withdrawn, subcooled, and sent to the middle of the upper column as feed. The remaining 465 mol are transformed into pure liquid nitrogen in heat exchange with the main condenser and sent to the top of the fractionation tower as liquid reflux. In the early 1920s, it was established that the 465 mol of reflux for every 1000 mol of process air was inefficient, because to produce oxygen even with a purity of 99.5%, only 200 mol of nitrogen reflux was required in theory. The balance 265 mol could be put to better use for the production of additional refrigeration. In actual practice, the nonessential nitrogen reflux was between 150 and 200 mol. There are several ways of exploiting the available excess nitrogen. One is to remove a portion of the uncondensed nitrogen from the high-pressure column, to recover part of its refrigeration capacity in a heat exchanger within the cold box, and then send it to the expansion machine for the generation of refrigeration. This refrigeration is recovered within the process system by sending the cold nitrogen through the primary exchangers in countercurrent with the incoming process air feed. The second alternative is to take a portion of the process air feed and send this stream directly to the upper (low pressure) column after first passing it through the expansion machine in order to reach the same operating temperature and pressure of the upper (low pressure) column. This latter procedure has been known as the Lachmann stream named after the developer.
ß 2006 by Taylor & Francis Group, LLC.
TABLE 3.5 Argon Content in Oxygen Product Oxygen
Argon
Nitrogen
95% 98% 99.5%
3.65% 2.00% 0.50%
1.35% Less than 0.10% Traces
Source: From Kerry, F.G., 2005. With permission.
The production of liquid nitrogen is a costly process in terms of energy because of its boiling point of 77.3 K at 101.3 kPa. As the purity of the oxygen product increases, so does the requirement for liquid nitrogen reflux. This increase in liquid nitrogen reflux is to remove argon from the oxygen product. Argon has a boiling point of 87.3 K, which is only 2.9 K different from that of oxygen. Therefore, to increase the purity of oxygen from 95% to 99.5% only the removal of argon enters the process of fractionation. This requires a relatively important quantity of expensive refrigeration. Table 3.5 may be of interest to the reader. Because most pyrometallurgical and chemical applications can operate economically with an oxygen purity of only 95%, it is not feasible to use a standard Claude cycle for the production of high-purity oxygen. For low-purity oxygen it is essential to come up with low investment and low energy consumption. For argon removal as well, there are some rigid specifications for many commercial combustion applications concerning the quantity of nitrogen in the oxygen product in various states in the United States. For example, in California the quantity of nitrogen in the oxygen to be used for combustion must not exceed 0.10% in order to keep the amount of oxides of nitrogen to a strict minimum in the effluent gas. To meet this standard the oxygen product must be kept at a minimum purity of 98.0% (Figure 3.15). Also, there exist specifications for the production of certain steels that limit any introduction of nitrogen in the smelting process to a hard-to-obtain minimum.
3.7 DISTILLATION EQUIPMENT The basic elements of a double-distillation system as it is known today are as follows:
3.7.1 UPPER (LOW PRESSURE) COLUMN The technology and theory of this has previously been described in Section 3.3.3.
3.7.2 LOWER (HIGH PRESSURE) COLUMN The importance of the lower column is probably the most misunderstood item by many practicing engineers. In fact, this column is particularly important because of the multiple duties it has to perform. Its design encompasses the following parameters: a. It must be designed for the maximum recovery of both oxygen and argon. b. The liquid nitrogen product condensed for the top of the column must be of the highest purity so that it may be used as reflux for the upper (low pressure) column, or as a marketable product. c. The amount of gaseous nitrogen to be liquefied in heat exchange with liquid oxygen in the main condenser must be sufficient to condense the nitrogen used as reflux, and to return 100% of the oxygen as well as argon to the sump of the tower.
ß 2006 by Taylor & Francis Group, LLC.
Oxygen nitrogen separation Theoretical calculation
5.0 4.0 3.0 2.0 1.0
Percentage of nitrogen in oxygen
Theoretical trays: Lower column 13; 19 practical Upper column 20; 28 practical
91%
92%
93%
94%
95%
96%
97%
98%
FIGURE 3.15 Oxygen–nitrogen separation—theoretical calculation. Theoratically the oxygen purity must exceed 96.5% to contain 0.10% nitrogen, the oxygen purity must reach close to 98% to attain less than 0.1% or less than 1000 ppm of nitrogen. The balance is argon. (ß Air Liquide, all rights reserved, 2006. With permission.)
d. It must be designed with an operating pressure low enough to save power, but high enough to drive the liquid nitrogen hydraulically to the top of the upper (low pressure) column for reflux. Actually, the pressure adjusts itself automatically. It keeps rising until the dew point is warm enough so that all the nitrogen gets liquefied. e. Sufficient distillation trays must be included in order to effect a 100% recovery of both the oxygen and argon products as stated above. f. The top of the high-pressure column should include a specially designed shelf to retain a portion of the high-purity liquid-nitrogen being condensed by the main condenser. It is this liquid nitrogen that is collected and used as reflux at the upper (low pressure) column.
3.7.3
MAIN CONDENSER
The main condenser, which has been called by many names such as ‘‘main vaporizer,’’ ‘‘vaporizer–condenser,’’ ‘‘reboiler,’’ etc., should rightfully be called simply ‘‘main condenser’’;
ß 2006 by Taylor & Francis Group, LLC.
because the direct purpose of the unit is to condense the incoming process air feed. The vaporization or reboiling of the liquid oxygen within the unit is the indirect result of condensation. This element is nothing more than a large storage receptacle of liquid oxygen descending from the upper fractionation tower. Regardless of its simple function, this element imposes very important design criteria for heat exchange and safety. The design should create conditions that ensure all internal heat exchange surfaces be wet with liquid oxygen at all times. If any dry spots occur, deposits of dangerous hydrocarbons such as acetylene, propane, and ethane will accumulate and may cause a catastrophic detonation. To avoid such potential accidents, it is important that a liquid oxygen recirculating system be created, either by the use of a liquid-oxygen pump, or by a thermosyphon. In either case the recirculating rate should be in excess of 100% of the evaporation rate. Some designs may not require a high recirculating rate, especially in the production and the extraction of large quantities of liquid oxygen, but such designs should be examined carefully before acceptance. Another requirement for the design of the main condenser is a high rate of heat transfer. The heat exchange between the liquid oxygen product on one side, and the incoming process air feed on the other can be affected by (a) High process-air-pressure in the (lower) high-pressure column, which means a high power-consumption at the main air-compressor (b) High heat-exchange-surface at the main condenser, which means high fabrication-cost for the unit (c) Use of a material with high heat-conductivity in fabricating the main condenser, which again means high investment cost (d) An optimum combination of all the above options Various designs have been tried out over the past several decades, but most air-separation plant designers have fairly well standardized on the use of extended surface aluminum exchanger cores that are available on a worldwide basis. These cores are used in a variety of combinations and packaged modules in order to provide a maximum of heat transfer with minimum capital cost. A differential temperature of 1.8 K has been recorded and is in operation since 1990. A recent report indicated that during heat transfer between the liquid-oxygen product and the incoming process in a main condenser the temperature differential was lower than 1 K, but this has not been confirmed officially.
3.7.4 LIQUID SUBCOOLERS Plants producing high-purity oxygen on a small scale are generally not equipped with liquid subcoolers. The availability of liquid reflux is large enough to avoid their use. In large air-separation units, however, because a fairly high quantity of gaseous nitrogen is withdrawn from the lower (high pressure) column and expanded in a machine to provide the required refrigeration capacity; an efficient liquid subcooler is a must for saving energy. Subcooling liquids prior to expansion through a J–T throttle valve avoids vaporization (mainly of nitrogen) and provides an increase of liquid reflux by approximately 8%.
3.7.5 PROCESS CONSIDERATIONS When the process air enters the (lower) high-pressure column and ascends to the top of the column, the main condenser condenses it at about 99 K. A portion of the liquefied air rich in nitrogen boils off, rises to the top, comes into contact with the main condenser that is at a temperature of 90 K or higher, recondenses, and descends once again, serving as liquid reflux
ß 2006 by Taylor & Francis Group, LLC.
for the ascending gaseous air flow. The oxygen and argon portions being less volatile are recondensed completely and accumulate in the sump of the tower. A portion of the liquefied nitrogen is used as reflux to make sure that all of the oxygen and argon are completely recovered. The liquid thus accumulated in the sump consists of nitrogen, oxygen, and argon, and is generally called ‘‘rich liquid’’ (liquid rich in oxygen). Care should be taken, however, that the oxygen content of this rich liquid be carefully calculated to arrive at an optimum balance within the designated operating pressure of the high-pressure column, and the equilibrium of the various components at that pressure. Assume that 1000 mol of air are divided into 550 mol of rich liquid with an oxygen content of approximately 40% and 450 mol of pure liquid nitrogen, which is retained on a shelf and serves as reflux after undergoing expansion to slightly above atmospheric pressure. This expansion across a throttle valve helps drop the temperature of the liquid to about 79 K (Figure 3.16). With regard to the heat-transfer capacity between the main condenser containing liquid oxygen and the ascending process air from the lower column, the following heat balance can be made: Assuming that 1 mol of air enters the lower (high pressure) column, 0.79 mol 6871 J/mol
0.21 mol 7252 J/mol
Unbalance differenial
5222.35 J/mol
0.45 2248 J/mol
1.0 mol 7252 J/mol
0.55 mol 1858 J/mol
FIGURE 3.16 Incoming process air ¼ 1 mol @ 7252 J=mol @ 506.67 kPa (5.06 bar) producing GOX þ GAN @ 6051 J=mol @ 101.3 kPa (1 bar) with an unbalance of 301 J=mol. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
Enthalpy of air at 506 kPa ¼ 7252 J=mol Enthalpy of rich liquid at 40% oxygen ¼ 1851 J=mol Enthalpy of liquid nitrogen ¼ 2248 J=mol 1 mol of air enters the lower column with 7252 J=mol and leaves with (1851 0.550) þ (2248 0.450) ¼ 2095.65 J=mol Therefore, the quantity of refrigeration being furnished by the liquid oxygen in the condenser to the lower column is equal to 72522095.65 ¼ 5222.35 J=mol, which is equal to the vaporization of liquid in the upper (low pressure) column. This energy is capable of vaporizing 5222.35=6830 (6830 J is the latent heat of vaporization per mol at 101.3 kPa) or 0.765 mol of gaseous oxygen of which 0.21 is sent to the primary- or front-end heat-exchangers as product, and the balance of 0.555% ascends the upper column for rectification purposes. The rich liquid at the sump of the lower (high pressure) column is withdrawn, subcooled, and sent to about the middle of the upper column for final rectification. As for the upper column, an enthalpy balance can be made for the 0.21 mol of oxygen and the 0.79 mol of nitrogen leaving the column: 0:21 7252 þ 0:79 6871 ¼ 6951 J=mol
(3:38)
This figure of 6951 J=mol is slightly below that of the incoming air feed, which has a total enthalpy of 7252 J. This imbalance of about 4% is made up by subcooling the rich liquid leaving the lower column, as well as cooling the liquid nitrogen reflux going to the top of the low pressure column, using outgoing gaseous waste nitrogen available from the top of the upper (low pressure) column as a refrigerant.
3.7.6 CRUDE-ARGON SEPARATION COLUMN Figure 3.17 shows the overall material balance treating 1000 mol of air including the separation of crude argon from the main distillation column.
3.8 DEVELOPMENT OF LOW OXYGEN-PURITY PROCESSES5 As noted, pyrometallurgical and many chemical operations did not require pure oxygen, and for many years air-separation designers strove to find an economically optimum solution for these industries. As plants grew larger in scale, however, so did the problem. Designers had to optimize the overall investment and operating costs for prepurification, cryogenic equipment, process air and product compression, as well as field costs for erection. Steel mills needed low-purity oxygen for smelting operations, but high-purity oxygen for oxy-cutting scrap and shape cutting for finished products. The product pressure for smelting had to be at 30 barG, and specified low nitrogen-content in the product oxygen to reduce nitriding of steel. Copper smelters needed only low-purity oxygen for smelting ore, but were willing to pay a premium for a product pressure of 3–4 barG. Syngas producers allowed a low purity for oxygen product, but at a pressure of 50 barG minimum. Since 1980, the increasing cost of oxygen compressors as well as the cost of field labor did not facilitate the problem. It is impossible to enumerate individually the variety of process cycles developed by imaginative process engineers, but a few merit some description. First of all, the question of oxygen purity has been quantified by industrial practice. For example, the steel industry smelts steel using the LD process with oxygen at 40 barG. Furthermore, although it can use oxygen with a low purity, the oxygen should be almost
ß 2006 by Taylor & Francis Group, LLC.
Nitrogen products N2 —780.76 O2— 2.04 Ar—1.30 784.10 mol
Nitrogen reflux N2 — 432.4 O2 — ppm Ar — ppm 432 mol Crude argon product N2 — 0.24 O2 — 0.16 Ar — 7 . 5 0 ( 9 5 % ) 7.90 mol
Oxygen product ( 99.7%) O2 — 207.4 N2 — ppm Ar — 0.6 208.0 mol Expansion turbine N2 — 84.5 O2 — 35.0 Ar — 0.5 120.0 mol Rich liquid Process air N2 — 781.0 O2 — 209.6 Ar — 9 . 4
N2 — 264.5 O2 — 174.6 (3.9%) Ar — 8 . 9 448.0 mol
1000.0 mol Approximate material balance Conventional air separation plant Nitrogen Product(s) Recovery—9.96% Oxygen Product Recovery—98.95% Argon Product Recovery—79.78%
FIGURE 3.17 Approximate material balance—conventional air separation plant. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
nitrogen free, which means that the minimum purity should be 98%, the balance being argon, as per Table 3.1. The copper industry can use a purity as low as 95% but at 3–4 barG.
3.8.1
THE LACHMANN PRINCIPLE (FIGURE 3.18)
As far back as 1902, Lachmann of Germany indicated that since impure oxygen would be more acceptable and economical to major industrial users, there was no point in wasting costly refrigeration to liquefy the total gaseous nitrogen for reflux necessary to produce high-purity oxygen. This concept was put into use in 1920 for plants producing liquid products that require extra refrigeration. The analysis of the Lachmann principle is simple enough. If y is the rising vapor and x is the descending liquid, therefore y ¼ 0:14 and x ¼ 0:40 (rich liquid) y ¼ x þ 0:315 (the nitrogen contained in the rich liquid)
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Waste nitrogen Turbine outage WP nitrogen HP air Gaseous oxygen MP air Main exchangers
Liquid oxygen to storage Liquid nitrogen to storage
LP column
FC
FC
Main vaporizer
Expansion turbines 2 x 100%
FC
HP column FC FC
LOX circulation pumps 2 x 100%
LOX filter LOX circulation pumps 2 x 100%
FIGURE 3.18 Use of nitrogen from the HP column to improve product recovery. (From Gomonet, E., Les Basses Temperatures et son Emplois, Librairie, J.B. Baillere et Fis, Paris, France, 1952. With permission.)
Hence the possible reflux in the upper (low pressure) column could be 0.17 instead of 0.47, which is required for high-purity oxygen. Therefore 30% of the nitrogen is available for liquefaction at 5 bar and could be used to supply extra refrigeration for other uses in the process cycle. For normal gaseous oxygen plants, this principle is carried out by taking 20%– 25% of the process air, passing it through the expansion machines at almost its dew point and inserting it into the upper (low pressure) column at a point just below the entrance of the rich liquid. This operation produces a low-purity product and saves energy.
3.8.2 THE OXYTON PROCESS5 In 1950, a paper was presented at the Physical Society in London introducing a new air separation process named Oxyton. It departed from the standard Claude cycle, because some effort was paid to the irreversibility of the rectification process of the lower (high pressure) column. The rectification was carried out at three pressure stages: 5.5, 2.5, and 1.3 barA. It included a third medium-pressure column (see Figure 3.19). This column received rich liquid from the lower (high pressure) column. After passing though a filter, the rich liquid was subcooled and expanded through a throttle valve before entering the middle tower. From there it flowed into an intermediate condenser–vaporizer built in the lower section of the high-pressure column providing refrigeration at 96 K, therefore offering a less irreversible separation to the rectification. The energy saved in the operation of the lower (high pressure) column was used at the middle-pressure tower to produce additional liquid nitrogen from the rich liquid at the top of the lower (high pressure) column. Thus, by improving the rectification of the latter tower, a greater quantity of nitrogen was withdrawn and utilized to provide additional refrigeration to the process. The power saved for the production of oxygen at 95% purity was significant for that era: 298 kwh=t. A plant with an oxygen-producing capacity of 300 t=d at 95% purity was designed and erected in northern Canada.
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94 K
% Nitrogen in the vapor
100 95
Claude HP column
90 85 80 Feed 75 50
99 K
60
80 70 % Nitrogen in the liquid
90
94 K
100
% Nitrogen in the vapor
100
95
Oxyton HP column
90
96 K
85 80
Feed
75 50
99 K
60
70 80 % Nitrogen in the liquid
90
100
FIGURE 3.19 Comparison between the Claude and Oxyton high-pressure columns. (ß Air Liquide, all rights reserved, 2006. With permission.)
3.8.2.1
Thermodynamic Analysis of the Oxyton Cycle
In the stripping section of the rectifying tower, the reflux ratio l is the ratio of molal reflux R and molal vapor flow V, or l ¼ R=V
(3:39)
The above equation is based on the assumption that the heavier product must be stripped down in the liquid in equilibrium, or Rx1 ¼ V y1
(3:40)
or lm ¼ y1=x1 (minimum reflux) Assume that F is the molal flow of a pure more-volatile component produced at the top of the tower; it can be said that V R ¼ F, therefore
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R(1=l 1) ¼ F
(3:41)
R ¼ [l=(1 l)] ¼ F
(3:42)
or
It can be said that the molal quantity of minimum reflux to obtain 1 mol of pure more-volatile product from a vapor containing y1% of the less volatile component may be expressed by the relation l=(1 l). If one assumes two separate stripping sections each operating at 5 barA with an infinite number of plates, and with a process air feed at its dew point of 99 K, the minimum reflux for the classical Claude and the Oxyton processes would be rm 99K ¼ lm=(1 lm )
(3:43)
where lm is the minimum reflux ratio ¼ y1=x1 ¼ 21=44 ¼ 0.447, resulting in rm ¼ 0:477=(1 0:477) ¼ 0:91 mol
(3:44)
In other words, for 1 mol of gaseous nitrogen extracted from the top of the fractionation column, one must obtain the above minimum quantity of liquid reflux at the feed. This quantity is produced by the main condenser–vaporizer at the nitrogen liquefaction temperature, which is 94 K at 5 barA. A less irreversible separation can be obtained, however, if an intermediate condenser operates at say 96 K. For example, lm 96 K ¼ y1=x1 ¼ 12:5=30 ¼ 0:417
(3:45)
rm 96 K ¼ lm 96 K=(1 lm 96 K) ¼ 0:715
(3:46)
which is lower than 0.91 mol of reflux necessary for the standard Claude cycle. The overall saving is 0.910.715 ¼ 0.195 mol, which is available for other purposes. The objective of this process was to narrow the area between the operating line and the equilibrium curve for the lower (high-pressure) column (see Figure 3.19). In this manner, the process was made more reversible. Although judged primitive when compared with modern systems, this cycle became the forerunner of other developments to reduce energy consumption and capital investment. Other more recent cycle developments using the same principle of a two-stage main condenser duty have evolved in the past few years. All of them, however, required extra cryogenic equipment and the inevitable increase in investment. With the expanding market for integrated coal gasification and nonferrous smelting applications, one cannot ignore or discount new process cycles that offer lower energy consumption combined with an attractive investment cost. 3.8.2.2 Oxyton Development In 1992, 52 years after the development of the original Oxyton, a large gas company in the United States sent out a request for a plant to generate 1989 t=d of contained oxygen in a 95% mixture, but at a pressure of 40 barA, and 85 t=d of pure nitrogen at 38 barA. The process engineers of Air Liquide remembered the original plant that had not been too successful in the market. They dug up the process drawings and went to work in improving the thermodynamic efficiency using contemporary ideas and equipment. The reversing regenerators for prepurification were replaced by molecular sieve adsorbers. Primary exchangers were
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replaced by nonreversing aluminum exchangers, and the third column, with its separate cooler, was replaced by a lean subcooler and a mixing column. The results of these modifications were spectacular. The oxygen product came out of the mixing column at the exact quantity, analysis, and at 4–5 barA, which eliminated the first casing of the oxygen compressor, thus saving a great deal of energy. The oxygen product generated was produced at the outlet of the process cycle almost at the same pressure as that of the inlet air. The nitrogen product came out as liquid nitrogen, which was pumped to the required pressure and vaporized in the main heat exchangers. No nitrogen compressor was needed, saving both capital, and operation and maintenance costs. Extra refrigeration capacity was obtained by taking 7% of the main air-process stream and upgrading its pressure using the expansion turbine as the driver. No extra booster air-compressor was required. If one applies only the consumption of energy by the main air-compressor to the process separation, the result is approximately 225 kw=t of product oxygen. This is very low by any standard.
3.8.3
VARIABLE-LOAD PLANTS
When direct oxygen-injection process of steel making took effect in the early 1950s, it reduced the time of smelting a heat of 100–200 t down to less than 1 h. The oxygen plant designers had to formulate an oxygen supply system operating on an on–off basis. At first this was achieved by designing a large steel tank operating at 30 barG with a system of piping and valves controlled automatically to inject oxygen at the prescribed pressure and timing as the vessels came into line. This system was acceptable but costly. Finally, a new air-separation process was developed, which did away with the high-pressure storage tanks. This resulted in a welcome saving for steel mills and for pyrometallurgical industries operating on a batch system. For maximum operating efficiency, air-separation plants have traditionally been designed to supply products at a constant rate. To accommodate a variable demand, a new process was designed achieving the same result within the cryogenic system, but eliminating costly external high-pressure tanks and ancillary equipment. This process has been in successful operation since 1975 not only for batch production (as in metallurgy), but also for continuously variable demand over 24 h periods (as in waste-water treatment). These processes operate between 30% and 100% of peak load demand. To make the unit cost-efficient, the cold box equipment is designed with larger inflow capacity than the upper limits of the process air compressors. Several designs are available on the market to achieve this variation in demand. Two, however, have a basic principle in common. Because the process cycles must vaporize an above-average demand, they include an extra condenser–vaporizer capacity as an additional piece of equipment, and the cycles include buffer (accumulator) tanks operating at cryogenic temperatures to store and supply extra liquid-oxygen product when needed. Moreover, they also include buffer (accumulator) tanks to store and supply either liquid nitrogen or liquefied air to the process when needed to maintain the refrigeration balance of the overall separation system. Incidentally, regardless of their basic small differences in design, a variable load oxygen process cycle can operate as described below in version A, or as in version B. 3.8.3.1
Version A
When oxygen demand is near the upper limit in one process cycle, liquid oxygen is pumped from the Liquid Oxygen (LOX) buffer tank to the sump of the upper (low pressure) column and vaporized in the main condenser–vaporizer. Meanwhile, the expansion machine will operate with a lower flow of nitrogen. The increased liquid oxygen in the sump will be vaporized by the additional nitrogen gas entering the main condenser–vaporizer from the
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lower (high pressure) column. This will increase Liquid Nitrogen (LIN) production, which will be sent to the LIN buffer tank for storage. When oxygen demand decreases LIN is pumped from the LIN buffer tank to the upper (low pressure) reflux plate, which will increase expansion machine refrigeration capacity, which in turn decreases vaporization capacity of the condenser– vaporizer. Excess LOX thus produced will be sent to the LOX buffer tank for storage. In other words, when the supplementary LOX demand is sent from the buffer tank to the condenser– vaporizer, the refrigeration from the condenser–vaporizer is used to liquefy extra gaseous nitrogen from the lower (high pressure) column, and when LOX demand is low, liquid nitrogen from the LIN buffer tank is sent to the cold box to maintain the designed refrigeration capacity of the process. During a 24 h period there is no gain or loss in liquid products. The only problem in this version is the fact that the buffer tanks are situated outside of the main cold box, thus requiring extra insulation, duplicate cryogenic pumps, and interconnecting piping for each buffer tank. This process cycle has been applied recently and very successfully in treating waste water at the west coast in the United States. In both Figure 3.20 and Figure 3.21 one can observe that the flow of oxygen follows the 24 h variation of the wastewater very closely and always at a slightly higher rate. Thus the reaction between the oxygen and wastewater is always complete. Four 24 h tests were carried out, and all tests performed as guaranteed. 3.8.3.2 Version B This alternative process cycle maintains the same basic principles as Version A, but has extra capacity for the main condenser–vaporizer, the LOX buffer tank and a liquid air (rich liquid) tank are supplied in a single but separate column. This extra tower is included in the main cold box, thus saving insulation and piping. The LIN buffer tank noted in version A has been replaced by a liquefied air (LAIR) buffer tank. Its contents are produced by taking a portion GOX production 325 305 285
Standard tons GOX
265
Curve 3 (Actual)
Curve 3 (Target)
245 225 205 185 165 145 125 14
15
16
17
18
19
20
21
22
23
24
1
2
3
4
5
6
7
8
9
10
11
12
13
14
Hours (02/28/02, 02:00 PM to 03/01/02, 02:00 PM)
FIGURE 3.20 GOX production. (Courtesy of Linde BOC Process Plants, 2006; CSDLAC. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
Gox production 440 420 400
Curve 4 (Target)
Curve 4 (Actual)
380 360 Standard gox
340 320 300 280 260 240 220 200 180 16 17 18 19 20 21 22 23 24 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 Hours (03/04/02, 04:00 PM to 03/05/02, 04:00 PM)
FIGURE 3.21 GOX production. (Courtesy of Linde BOC Process Plants, 2006; CSDLAC. With permission.)
of the main process air, boosting it to a higher pressure, precooling it, and then passing it through the auxiliary vaporizer. In this auxiliary condenser–vaporizer, liquid-oxygen product is vaporized at the prescribed operating pressure and exits the plant after passing through the main heat exchangers. In turn, process air is partially liquefied, subcooled, and sent to the liquefied air buffer tank. By the time it enters its buffer tank, one may call it rich in oxygen or rich liquid as it is known in normal air plants. It joins the rest of the precooled main process air in the bottom of the lower (high pressure) column. The LOX product is sent to the condenser by one of two parallel 100% capacity cryogenic transfer pumps (see Figure 3.21 and Figure 3.22). The former figure shows that during a single blow of oxygen the level of the oxygen tank drops quickly, whereas during a 24 h period, as shown in the latter figure, there is a series of oxygen blows and a subsequent series of level variations in the oxygen accumulator tank. The blows were carried out both in the reverbatory and convertor separately (see Figure 3.22 and Figure 3.23).
3.9
EXERGY
Although the term exergy was used by Trepp in 19626, and by Patela in 1964 in a few individual cases, the word was widely accepted by the industry after a general paper was published by Robert Evans of Georgia Institute of Technology in 1966. The second law shows the influence of irreversibilities within system components, and can provide methods to minimize the losses of energy and capital. Even in 1939, Kapitza concluded that conventional isentropic efficiencies were not a true measure of low-temperature applications when processes were at different temperature levels. In 1979 and thereafter, engineers began to use the exergy method for analyzing the heat exchanger and turboexpander performance. Exergy has been widely used in Europe and in
ß 2006 by Taylor & Francis Group, LLC.
Tiempo real
Nallsis historic 11:17:37 1h
4h
Jul 24 11:17:37
Tiempo real
2 h 15 min
Zoom in Min Jul 24 11:51:22
Zoom out 30 min
Jul 24 12:25:07
Valores 13:32:37
Unavail
96.6
LocsaNivelEst95
10 min
Jul 24 12:58:52
AIT−2040
Jul 24 13:32:37
Unavail
56.5
FT-2040
100
110.0
Unavail
98 96 94
PT-2040
92
35.0
Unavail
90 88 86 84 82 80 11:17:37
Imprimir 13:32:32
FIGURE 3.22 Variable Oxygen used in stop start operation for converter and reverbatory furnaces covering 1 hour and 40 minutes. (ß Air Liquide, all rights reserved, 2006. With permission.)
Analisis
Nalisis historic 23:55:00
Zoom in 1h
4h
Jul 24 23:55:00
Min Jul 25 01:28:45
Valores
Tiempo real
6 h 15 min
Zoom out 30 min
Jul 25 03:02:30
Jul 25 04:36:15
06:10:00
AIT-2040 Unavail
96.6
LocsaNivelEst95
10 min Jul 25 06:10:00
79.6
64.9
FT-2040
180
138.4
162
168.5
144 126
PT-2040
108
34.2
32.4
90 72 54 36 18 0 23:55:00
Imprimir 06:09:48
FIGURE 3.23 Variable Oxygen used in stop start operation for converter and reverbatory furnaces covering 6 hours and 10 minutes. (ß Air Liquide, all rights reserved, 2006. With permission.)
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Russia for maximum available work function. More recent work on the second-law application in cryogenics has been aimed at the analysis of component parts rather than in systems. A detailed second law–system analysis of energy availability improvement, the magnitude, location, and cause of each significant irreversible loss in the system must be determined. Losses may include mass and heat gains (external and internal), flow frictional losses, and entropy production from flow stream mixing. D.W. Townsend7 of ICI Petrochemicals, Process Division, clarified the term exergy lucidly in 1979. He stated that the application of exergy is the best single concept to carry out an energy-efficiency study. It is defined as the maximum work that can be derived from a reversible system, or Ex ¼ (H H0 ) T0 (S S0 )
(3:47)
A process depends specifically on its cooling-medium temperature T0 and by integrating it can be shown that DS ¼ DH{[ ln (T2=T1 )]=(T2 T1 )}
(3:48)
Linnhoff8 has shown that Equation 3.48 is an excellent approximation even when linear TDH profiles are involved. If gaseous pressure changes occur, the ideal gas assumption holds: DS ¼ [(Cp dT)=T] [(RDP)=P]
(3:49)
DS ¼ DH{[ ln (T2=T1 )]=T2=T1 } R ln (P2=P1 )
(3:50)
and it follows that
Equation 3.49 and Equation 3.50 have been accurate enough for most applications. An irreversible process is one in which the potential for doing work is lost and it can be shown that X Ex(in) ¼ Ex(out) þ Ex(irreversible) (3:51) This equation offers an objective basis for the evaluation of the second law. Exergy thus provides a useful tool for the investigation of efficiencies to reduce energy requirements in refrigeration systems, which include compressors, condensers, system exchangers, flash separators, and general process exchangers; to reduce energy concepts; and at minimum capital investments. Process and project engineers should strive to meet this target despite the poor Carnot efficiencies encountered in cryogenics. More recently exergy or the analysis of the second law has already been used to select optimal LNG process cycles employing front-end propane refrigeration, and to evaluate cryogenic versus noncryogenic processes for the production of pure nitrogen.
REFERENCES 1. King, G.R. 1971. Modern refrigeration practice, 48–61. New York: McGraw Hill. 2. Timmerhaus, K.D. and Flynn, D. 1989. Cryogenic process engineering, 291. New York: Plenum Press. 3. Barron, R.F. 1971. Cryogenic systems (Monogram on cryogenics), 2nd ed., 151–200. New York: Oxford University Press. 4. Garcia, L. 1995. Basic principles of air separation. In Practical operating manual. Paris: L’Air Liquide.
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5. Grunberg, J.F. 1950. The oxyton cycle in the production of tonnage oxygen. In The Industrial Chemist. London: The Physical Society. 6. Trepp, C. 1962. Adv Cryog Eng 7. 7. Townsend, D.W.T. 1980. Second law analysis in practice. Chem Eng (October): 628–633.
ß 2006 by Taylor & Francis Group, LLC.
ß 2006 by Taylor & Francis Group, LLC.
4
Rare (Noble) Gases
The rare gases, sometimes referred to as noble gases, are situated in the zero group in the periodic table of elements between the most electronegative (the halogens) and the most electropositive (the alkali metals) elements. Because their electron shells are completely filled, they are inert, that is, do not normally form chemical compounds. Their ionization energy, the energy required to remove an electron from the valence shell, is greater than for any other element, and conversely, their electron affinity, the energy released in taking up another electron, is zero. Nevertheless, it has been found that for the heavier inert gases, whose outer electrons are farther removed from the nucleus and whose force of retention is weaker, and electrons can indeed sometimes be removed by strong electromagnetic fields. Fluorides of krypton, xenon, and radon have thus been prepared, and recently krypton has actually been inserted into acetylene mainly to form an ionic bond with carbon. (One may thus say that a more appropriate classification in the periodic table would be in group VIIIA.) (Table 4.1). During the early part of the twentieth century most interest in the rare gases was in the research laboratory. It was oxygen that took the center stage. In the latter part of the century, however, the industrial uses of rare gases grew by leaps and bounds in applications ranging from welding to deep-sea diving to medical imaging to industrial and commercial lighting.
4.1 HELIUM Helium was discovered almost simultaneously in 1868 by Pierre Janssen of France and Norman Lockyer of England when each conducted a spectrographic examination of the solar chromosphere during an eclipse in India in 1868. Their findings were presented at the same meeting of the Royal Society, but Lockyer has been granted precedence because he had invented the instrument that Janssen had also used. They had observed a line similar to but not quite the same as the D1 and D2 lines of sodium. It was therefore labeled D3. When it was later recognized that due to a new element the latter was named helium (Helios in Greek). In 1895, when Sir William Ramsay spectroscopically analyzed a gas given off by heating clevite, a uranium-bearing ore, he saw the identical D3 line. In 1903, both Ramsay and Frederic Soddy determined that helium was a product of the disintegration of radioactive minerals. When the concept of thermonuclear fusion was developed during the 1940s it was realized that this process, fusing hydrogen into helium, was responsible for the energy emitted by the sun and all other stars. Although helium accounts for 23% of the known mass of the universe it occurs in the Earth’s atmosphere only at 5 vppm since the Earth’s gravity is not strong enough to prevent its escape to outer space. Helium is a colorless, odorless, and tasteless gas. It can be liquified at 4.224 K and may be solidified, but only at a pressure of at least 26 barA. At any lower pressure it remains as a liquid, even at lower temperatures. As a liquid its index of refraction is very close to that of its gas so that it is very difficult to observe it in a flask. It has six known isotopes, but only two are stable, that is, helium 4 and helium 3. The other four isotopes are radioactive and decay
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TABLE 4.1 Physical Properties of (Rare) Noble Gases Normal
Element
Atomic No.
Molecular Mass (kg/kmol)
BPt K (K at 1.103 bara)
Vapor Pressure (mbarb)
Latent Heat (MJ/kmolc)
Specific Heat (kJ/kmol Kd)
Helium Neon Argon Krypton Xenon Radon
2 10 18 36 54 86
4.0026 20.09 39.948 83.80 131.30 222.0
4.224 27.09 87.28 119.8 164.1 211.2
51 433 688 731 816
0.083 1.79 6.65 9.04 12.63
20.80 20.79 20.78 20.87 21.02
a
Normal point is at a pressure of 101.325 kPa. (Atmospheric pressure at sea level.) Vapor pressure is at the triple point. c Latent heat of vaporization is at normal boiling point 101.325 kPa. d Molar specific heat capacity at constant pressure in range from 0 to 1 bar at 298.15 K. Source: From Pauling, L. in General Chemistry, Periodic System of Elements, W.H. Freeman & Company, San Francisco, 1958. With permission. b
very quickly. Helium 4 is the most abundant of the isotopes. It is present in the Earth’s atmosphere about 700 times the concentration of helium 3. What helium 3 does exist in the air is due to the b-decay of tritium, hydrogen 3. Timmerhaus and Flynn1 have stated that when liquid helium 4 is subcooled below 2.17 K it splits up into two phases, that is, helium 1 and helium 2. At lower temperatures, the former has a viscosity of 3 106 Pa s; and the latter phase exhibits properties of superfluidity, inasmuch as its viscosity and resistance to flow almost vanishes. The viscosity of helium 2, at 1.3 K, is approximately 3 107 Pa s, or even lower (1012) in narrow channels, and its thermal conductivity reaches 86,500 W=(m K) that is 1000 times greater than that of copper.
4.1.1
SOURCES
OF
HELIUM
Originally, helium was recovered by heating monazite earth to a temperature of 1000 K. Shortly afterward, and with the industrial development of air separation plants, Georges Claude2 in 1908, suggested that it was possible to collect an uncondensed mixture of neon– helium from the top of the main condenser (MC) of the distillation section of an air separation plant. If this mixture was allowed to accumulate it would lower the efficiency of heat transfer of the MC, and therefore lower the productivity of the air separation plant. When syphoned off, this mixture could then be cooled to the temperature of liquid hydrogen, and when passed through a bed of activated charcoal, the neon could be adsorbed in a solid state, and the helium freed as a gas. This latter process was still used up to World War I. But although atmospheric air contains 5.3 vpm of helium its recovery is uneconomical, except for small laboratory work. Soon after World War I a natural gas well at Excell, Texas, in the Texas Panhandle was discovered to contain a fair quantity (0.70%) of helium. It was also shown that it was more economical to extract this helium for industrial uses by using a cryogenic process. Consequently, the U.S. Government told the Bureau of Mines to set up an office at Amarillo Texas, to take charge and develop the recovery and conservation of helium. From 1918 to 1961, the Bureau of Mines had fairly tight control over the use of helium, and its consumption was restricted to military operations, the inflation of weather balloons, and lighter-than-airships.
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Natural gas processing
Helium processing
Helium distribution
Natural gas plant End-use customers
High-purity helium (gas)
Fuel gas
Gas distributors
Crude helium Gas tube trailer
(50% -70% He) Upgrader
Cylinders for end-use customers
Liquid storage
PSA
End-use customers Liquefier
Bulk liquid helium transport
Bulk distribution centers
Tube trailers, cylinders, dewars for end-use customers
Natural gas reservoir
FIGURE 4.1 Recovery of helium from natural gas reservoir to consumer in gaseous and liquid form. (Courtesy of Praxair, Inc., 2006. With permission.)
The Bureau’s other helium-bearing natural gases were located nearby at Keyes, Oklahoma; Shiprock, New Mexico; and Otis, Kansas. After 1960, private industry entered the helium recovery field by long-term contracts with the U.S. Government. In 1961, a privately owned plant started up in eastern Arizona with an annual production of 1.875 MM Nm3. Then, in 1966, another plant was built in Otis, Kansas, with an annual production of 4.019 MM Nm3. And in 1966, a third plant was erected at Elkhart, Kansas, with an annual production of 3.750 MM Nm3. These plants produced unrefined helium of 50%–85%, which was then purchased by the U.S. Government and stored underground in a natural gas deposit for future use (Figure 4.1). On a worldwide basis, the following sources are producing helium in the year 2000 (Table 4.2).
4.1.2 GENERAL PRINCIPLES
OF
RECOVERY
OF
HELIUM
Helium was early discovered in very limited quantities in radioactive mineral springs at Bath, England. However, not much work was carried out on its physical properties. It was at the turn of the twentieth century that helium was also discovered in monazite sands,
TABLE 4.2 Worldwide Sources of Helium Source Kansas Texas Panhandle Poland Holland Germany North Sea, Leman Bank France
He (%)
CH4 (%)
N2 (%)
C2 Plus
2.0 0.7 0.4 0.045 0.04 0.03 0.001
— 73.2 56.01 81.3 42.5 94.7 97.1
23.0 14.3 42.75 14.35 56.5 1.3 0.3
Balance Balance Balance Balance Balance Balance Balance
In Canada, Algeria at Arzew and Skikda, and Romania there are also small sources. Source: Courtesy of CryoGas International Publication, 2002. With permission.
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TABLE 4.3 Typical Analysis of Uncondensed Stream from Main Condenser of Air Separation Unit Nitrogen Neon Helium
70.0% 23.0% 7.0%
Source: From Ruhemann, M. in The Separation of Gases, Oxford University Press, London, 1945. With permission.
phosphates-containing rare earth metals and thorium, at gravel pits in North Carolina, from which helium could be extracted, albeit with difficulty. After this important discovery, serious and extensive experiments were carried out on helium by Kamerlingh Onnes in Holland on the element’s physical properties, both in gaseous and liquid phases. When industrial air separation plants came into practical operation, Georges Claude suggested that helium could be recovered from the MC where both helium and neon accumulated in an uncondensed vapor. Unless removed, this created inefficiency in the heat transfer at the MC. The approximate analysis of the takeoff gas in the uncondensed stream was as follows3 (Table 4.3). The nitrogen was easily separated by compressing the mixture to 50 bar and by passing the gas through a reflux condenser that has been filled with liquid nitrogen. The remaining traces of nitrogen were removed by passing the effluent through activated charcoal, or over hot calcium, or magnesium in order to save the neon. To separate neon from helium, the use of nitrogen as a refrigerant was out of the question because its triple point is 63 K, and at this temperature its vapor pressure is 433 mbar, which is very difficult to maintain industrially due to the necessary use of a deep vacuum operation. Nevertheless, Meissner4 succeeded in separating the two elements with the use of liquid hydrogen as a refrigerant. This was the process used up to 1918, at the Physikalische—Technische Reichsanstalt at Charlottenburg in Germany. With the discovery of helium in natural gas at the Texas Panhandle the recovery of helium became a large-scale industrial process, and the Bureau of Mines soon found that for large flows of natural gas a helium content of less than 0.1% was not viable for production. In fact, the recovery of helium is of little interest when its content is less than 0.5%, at least in the United States. For a content of 0.5% and up to 2.0%, recovery is economically viable. Nevertheless, in former Soviet Union, the engineers were exploiting natural gas with only 0.1%–0.3% of helium, no doubt, for its military potential. The presence of significant nitrogen in natural gas more often than not indicates that helium is also present. Moreover, the presence of nitrogen is of definite advantage because it can be used as a refrigerant cycle after a methane cycle. The resulting liquefied nitrogen can be used to augment the recovery as well as the purity of the helium. In short, the combined use of both nitrogen and methane evolves into a classic cascade cycle, which is the most economical, especially if the natural gas also contains ethane and propane. Depending on the composition of the natural gas, a wide variety of process cycles have been designed, most of them operating on multistage refrigeration cascade process cycles. The first effect is a ternary mixture of methane, nitrogen, and helium. Methane is the first to be liquefied. As the solubility of helium in liquid methane is insignificant the helium remains in the gas. The binary mixture of nitrogen and helium remaining, however, is difficult to separate, irrespective of the nitrogen content. If one considers the binary system of helium and nitrogen, the latter being the main refrigerant, though the boiling points are far apart and the vapor–liquid equilibrium curves are well known, their separation is not easy, depending on the purity desired for the helium.
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TABLE 4.4 Helium–Nitrogen Mixture P (bar) 10 20 40 60 100
% He in Vapor % at 80 K
At Equilibrium % at 70 K
79 91 95.5 97.2 98.5
93 98 99.2 99.6
Source: From Ruhemann, M. in The Separation of Gases, Oxford University Press, London, 1945. With permission.
For very high purities, of say 99.5%–99.85%, very high pressures are needed to attain high concentrations of helium vapors in equilibrium with liquids at 80 and 70 K. From Table 4.4 it is obvious that with nitrogen as a refrigerant, one can only hope to reach a refrigeration temperature of 77.3 K, the atmospheric boiling point of nitrogen. Therefore, one has to consider either the use of a higher pressure or the use of liquid nitrogen boiling at a lower pressure. Presently, it is common to use a reasonable pressure of around 40 bar for the closed nitrogen cycle and a much higher pressure for the helium cycle. The latter combination will result in an optimum consumption of energy for the final stripping of nitrogen. This concept is especially applicable for helium-bearing natural gases where the helium is present in a comparatively small quantity and nearly total recovery is mandatory (Figure 4.2).
N2 liquefier Crude He Crude He
Crude He separator He purifier Crude He
Heat exchanger
He heat exchanger
N2 gas
He heat exchanger We
Natural gas Compressed natural gas Natural gas
95% N2
Pure He separator N2
N2 + He in solution Separator
98.5% He
FIGURE 4.2 Helium separation by Bureau of Mines. (From Timmerhaus, K.D. and Flynn, T.M., Cryogenic Process Engineering, Springer Science & Business Media, 2005. With permission.)
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4.1.3 RECOVERY PROCESSES FROM NATURAL GASES Before applying a cryogenic refrigeration cycle to any natural gas the entire gas stream has to be treated in order to remove all the cryogenic undesirables such as carbon dioxide, hydrogen sulfide, and water, which may be present in the process gas. If carbon dioxide content is above 2.9% by volume one must use either an amine scrubbing system, or refrigerated methanol scrubbing (Rectisol). A further treatment with a molecular sieve containing activated alumina upstream is then needed to make sure that the prepurification treatment operates at maximum efficiency. If carbon dioxide is less than 2.9%, the pretreatment can be, and has been carried out, by the exclusive use of molecular sieve adsorption. The refrigeration process may be considered as operating in two stages. In the first stage, the extraction involves the removal of both the methane and a major portion of the nitrogen by liquefaction. If too much of the nitrogen is removed, however, there may be a high loss of precious helium, because of the latter’s solubility in liquid nitrogen. If any small quantities of nitrogen and helium are left they will degrade the natural gas in terms of valuable calorific value. The remaining binary mixture of nitrogen and helium involves the denitrogenation of helium. This part of the overall process uses a standard Claude closed nitrogen cycle with a compressor and an expansion turbine. Depending on the final composition of the helium required it may also require the use of a high-pressure booster compressor for the helium. This high compression is also useful for transporting the helium product in cylinders elsewhere for further purification. Typical examples are from the Bureau of Mines (see Figure 4.2)6. The natural gas is compressed to a pressure of 42 bar, treated for the removal of carbon dioxide, hydrogen sulfide, and water, then sent through a primary heat exchanger (PHX), where it is cooled by the returning stream of cold natural gas containing some nitrogen, but free of helium. The main process stream continues through a throttle valve and is expanded into an MC at a pressure of 18 bar and a temperature of 125 K. In the MC all hydrocarbons are condensed and returned to the PHX, where they serve to cooldown the incoming natural gas. The MC is also refrigerated by an incoming stream of cold gaseous nitrogen at practically atmospheric pressure from the expansion turbine. Once this nitrogen gives up its refrigeration it is then returned to the inlet of the nitrogen-cycle compressor. The vapor at the top of the condensed natural gas in the MC consisting of approximately 60% helium, 40% nitrogen, and traces of methane, is extracted and sent to the inlet of the helium compressor after giving up its refrigeration in a helium heat exchanger (HHX). The principal refrigeration system is a standard Claude cycle employing a closed nitrogen cycle with a compressor boosting the gaseous nitrogen to a pressure of 42.6 bar. After passing through a nitrogen and nitrogen heat exchanger (NHX) where it is cooled by a returning stream of low-pressure nitrogen, it is split into two streams. One stream enters an expansion machine and is sent as a cooling gas to the top of the MC. The second stream is cooled by the expanded nitrogen stream, and is throttled as a liquid and is sent to the condenser–evaporator (CE), where it serves to strip the nitrogen from the helium mixture. It also upgrades the final product helium to a purity of 98.5% by maintaining a low temperature at an activated charcoal filter, as the gaseous product helium passes through it. At the bottom of the CE, the liquid nitrogen is removed, throttled to a lower temperature, and enters a separator (S). At the separator more of the nitrogen is condensed, whereas the uncondensed helium vapor containing progressively less nitrogen is removed and sent to the C-1 condenser to be recycled into the process for further stripping and upgrading. In order to achieve a higher purity, of say 99.995% or even up to 99.9995 þ %, it is necessary to treat the 98.5% helium separately using liquid nitrogen boiling at subatmospheric pressure (20 kPa), and removing the nitrogen with activated carbon. The vessels containing the activated carbon are also kept cold by encasing them in a liquid nitrogen cooled jacket.
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4.1.4 APPLICATIONS
OF
HELIUM7
Because helium is chemically inert, nonflammable, has a high thermal conductivity, a low molecular weight, a small molecular size, and is nontoxic, it has many applications. It has a high arc temperature and a high heat transfer rate. This makes it ideal for its use in inert gas shielding, or plasma gas welding of steel, aluminum copper, nickel, stainless steel, titanium, and many other alloys. It is also beneficial for the smelting and refining of advanced materials such as niobium, tantalum, zirconium, and titanium. Having a very small molecule, helium’s high penetrating quality is very advantageous in leak detection, especially in aerospace equipment, as well as in industries dealing with semiconductors, and nuclear, cryogenic vacuum, refrigeration, and food canning equipment. Because of its low solubility and low boiling point (4.2 K), as well as its nonflammability even at very low temperatures, helium is employed as a purging and pressurizing agent for liquid hydrogen fuel systems of rockets and spacecraft. Helium is very useful in deep-sea diving and work in underwater caissons, replacing nitrogen in admixture with oxygen. Breathing it under high pressure does not result in the rapture of the deep, a dangerous narcosis caused by nitrogen, nor does it cause agonizing bends when a diver resurfaces rapidly. Helium–oxygen breathing mixtures have opened the door wide to undersea scientific exploration and made possible the development of offshore oil and gas resources even though the mixtures have the minor disadvantage of raising the pitch of divers’ voices. Helium is critical for producing optical fibers to be used in telecommunication cables. It also allows manufacturers to grow crystals for electronic semiconductors or crystals with optical properties needed to produce masers and lasers. Apart from applications of gaseous helium, liquid helium has developed as one of the most important new materials of science. It helps to cool certain compounds below a critical temperature, at which they become superconductors. Superconducting magnets are used in magnetic resonance imaging (MRI) that has become a powerful tool in diagnostic medicine. The application of a similar technology, nuclear magnetic resonance (NMR), is also invaluable in probing the structure and properties of chemical compounds. Liquid helium also finds application in power generation and transmission, fusion research, magnetohydrodynamic (MHD) propulsion of ships, superconducting magnetic energy storage (SMES), electronics, and magnetically levitated (Maglev) transportation systems. In order to give an idea of the year 2000 consumption of helium in the United States the following approximate figures in Table 4.5 are of particular interest8.
TABLE 4.5 Annual Consumption in the United States Industrial and scientific MRI Welding Lift gas Fiber optics Purge and pressure gas Leak detection Controlled atmospheres Breathing gases Other
18% 18% 16% 14% 9% 8% 6% 3% 3% 5%
Source: Courtesy of CryoGas International Publication, 2002. With permission.
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FIGURE 4.3 Helium recovery after use. (Courtesy of Praxair, Inc., 2006. With permission.)
4.1.5 CONSERVATION
OF
HELIUM
In view of the simple fact that helium drifts off into outer space and is lost forever, every effort should be made by industry and all users to recover helium at every opportunity. The U.S. Government has already taken the lead in storing helium within natural gas storage caverns. This idea of conservation should be followed by ultimate users. This precious element should not be wasted for toy balloons and parade floats. Equipment has already been designed to recover helium at major users (Figure 4.3).
4.2 NEON 4.2.1 GENERAL Neon, which derives its name from the Greek (meaning new), has an atomic weight of 20.183, with three stable isotopes of 20, 21, 22, and is found in the atmosphere at 15.385 106 vppm. It is inert, but in certain conditions it can form a compound with fluorine. Liquid neon is clear and colorless, and has a boiling point of 27.09 K, a triple point of 24.56 K, and a vapor pressure of 433 mbar, at its triple point. The electrical discharge of gaseous neon at ordinary voltages and currents gives off an intense red.
4.2.2 SOURCES
OF
NEON
Neon has been extracted from the atmosphere as a by-product in air separation plants. As described in Section 4.1.2, neon can be recovered from the noncondensable stream removed from the MC of the air separation column. After catalytic oxidation to remove all traces of any hydrogen, and further treatment in an additional column and condenser assembly, the analysis is improved to the following9 (Table 4.6).
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TABLE 4.6 Typical Analysis of Neon Stream after Complete Hydrogen Removal Nitrogen Neon Helium
70% 23% 7%
Source: From Ruhemann, M. in The Separation of Gases, Oxford University Press, London, 1945. With permission.
4.2.3 RECOVERY
OF
NEON
In the treatment of this last analysis it is important to remove the nitrogen completely, as any nitrogen left may spoil the color when neon is used for lighting. By passing the mixture at a pressure of around 50 bar through a reflux exchanger surrounded by a bath of liquid nitrogen, it is fairly easy to remove the bulk of nitrogen. The remaining traces can be adsorbed by a bed of activated charcoal. If the latter material also causes a high loss of neon, then one must consider the use of either hot calcium or magnesium. It should be noted, however, that the separation of neon from helium is very difficult as well as expensive. The use of nitrogen as a refrigerant is out of question because the lowest temperature attainable even as a liquid at its triple point is 63.2 K, and its vapor pressure at this temperature is only 126 mbar, which is difficult to maintain industrially due to the necessity of a continuous deep vacuum operation. On the other hand, the critical temperature of neon is 44.41 K and that of helium is 5.2 K. One, therefore, had to consider the use of liquid hydrogen. Alternatively, one may use Sterling cycle refrigerators to provide low-temperature cooling. Meissner10 succeeded in separating a neon–helium mixture in a fairly large scale by using liquid hydrogen in the early 1930s. This treatment, however, presents a problem because the boiling point of hydrogen is 20.4 K and that of neon is 27.2 K. The triple point of neon on the other hand is 24.6 K, which is above the boiling point of hydrogen. Care has to be exercised, therefore, to make sure that the neon does not solidify and block the tubes. To avoid this problem, one has to maintain the operating pressure above the liquid hydrogen at 3 bar, which is sufficient to raise its boiling point above the triple point of neon. Another method is to raise the temperature differential between the condensing neon and the evaporating hydrogen by reducing heat transfer surface in equipment.
4.2.4 INDUSTRIAL RECOVERY
OF
NEON
In recovering neon, one must consider the final purity desired, especially in terms of helium content. Helium is not really an undesirable impurity if the mixture is to be used in fluorescent lighting fixtures as a certain quantity of helium adds to the effect of the final color. If on the other hand, a high purity of neon is the objective (especially if the final product is liquid neon), then a two-phase system must be designed. The first phase (Table 4.6). The noncondensable mixture of helium, neon, nitrogen, and some traces of hydrogen is withdrawn from the top dome of the MC and led into the bottom of a separate column equipped with a reflux condenser, where the vapors are led through a bath of liquid nitrogen. The condensibles, mostly liquid nitrogen, fall to the bottom and are returned to the MC of the air separation unit. The remaining noncondensables, a mixture of gaseous neon–helium, are withdrawn from the top dome of the condensing unit and led into a second phase of purification involving liquid hydrogen.
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In this secondary and final phase, the mixture from the auxiliary reflux condenser is mixed with a small amount of oxygen and fed into a catalytic burner to remove any traces of hydrogen. After the stream passes through an after-cooler and dryer, the final mixture of neon–helium enters an adsorption trap encased in, and cooled with liquid nitrogen. Finally, the resultant stream passes through a closed cycle hydrogen-refrigerated exchanger and enters a secondary reflux exchanger where it is liquefied in a bath of liquid hydrogen.
4.2.5
INDUSTRIAL APPLICATIONS
OF
NEON
Apart from its use in fluorescent light fixtures, neon has also found application in high-energy physics. According to Isalski11, the study of subatomic particles involves the use of liquid hydrogen-filled bubble chambers with a target residing at their center. When highly accelerated atoms strike the target, reactions may occur, releasing subatomic particles. Liquid hydrogen has the property of creating tracks that can be photographed when a subatomic particle is released. Liquid neon is more dense than hydrogen and has been used outside the hydrogen chamber in order to slow down the reaction, resulting in a slow-motion effect. Liquid neon also has an advantage over liquid helium as a low-level refrigerant. It is less expensive per unit heat of vaporization, and a precise temperature control can be easily achieved by varying the saturation pressure of liquid neon. Liquid neon is also used by the military in infrared detection equipment to improve light sensitivity. It actually reduces molecular motion within the equipment.
4.3 4.3.1
ARGON GENERAL
Argon, which derives its name from the Greek (meaning late comer), has an atomic weight of 39.948, and three isotopes 36, 38, 40, and is found in the atmosphere at 0.934%. The liquid phase is a clear colorless fluid with a normal atmospheric boiling point of 87.28 K at 101.325 kPa. Its triple point is 83.8 K, its critical pressure is 4860 kPa (48.6 bar), and its vapor pressure at its triple point is 688 mbar.
4.3.2
SOURCES
OF
ARGON
In 1785, H. Cavendish determined experimentally that about 5% of the so-called dephlogisticated air had different physical characteristics than ordinary air. In 1884, Ramsay and Rayleigh discovered that the atomic weight of atmospheric nitrogen was heavier than that of nitrogen produced by chemical reaction. The difference, they found, was due to a new inert gas in the atmosphere, which they have named argon. Further studies proved that this new element formed 0.934% by volume of the atmosphere, or 1.25% by actual weight. It also proved to be a very valuable factor in our modern industry.
4.3.3
RECOVERY
OF
ARGON
The atmospheric boiling point of argon (87.3 K) is closer to that of oxygen (90.2 K) than to that of nitrogen (77.3 K). As can be seen from its physical data, argon has a boiling point (87.3 K) between those of oxygen and nitrogen, namely, 90.2 and 77.3 K, respectively. In fact, it is closer to that of oxygen (90.2 K). In other words, a high recovery of argon implies a complete separation of oxygen and nitrogen, which requires extra energy and equipment. Forty to 50 years ago, a recovery of 40% of argon from a standard air separation plant was considered high for any definite guarantee. Today, however, designers are committing
ß 2006 by Taylor & Francis Group, LLC.
themselves to recoveries as high as 95% in certain conditions of operation, especially in liquid plants. Optimum recovery of argon depends on three factors: (a) Increasing the number of trays in the upper (low pressure) column. This factor, however, increases the investment because of extra trays, longer piping, extra outer casing, and insulation. There is also an increase in overall pressure drop, which increases energy consumption. (b) Increasing efficiency of the distillation trays. This raises the recovery, but only marginally. (c) Increasing the total quantity of liquid nitrogen production, hence increasing the total nitrogen reflux both at the lower and upper columns. In this manner the argon will not escape in either the product or waste nitrogen streams. This option, however, requires extra refrigeration capacity with the use of an external nitrogen recycle compressor. This indicates a higher energy consumption. A combination of all the three factors can provide a good balance between energy and investment. One should note that there is a definite relationship between the quantity of nitrogen withdrawn from the lower column of an air separation unit, whether in gaseous or liquid phase, and the recovery of argon. The quantity of nitrogen available for reflux may be limited by (a) The amount of nitrogen, or nitrogen–oxygen mixture, expended for refrigeration in the expansion turbine (b) The amount of nitrogen withdrawn as a product either as a liquid or as a gas In other words, recovery of argon (RA) is proportional to N2 and processed air (PA). To overcome this above-mentioned limitation and to increase argon recovery to its maximum, it is mandatory to produce extra refrigeration by an external cycle using nitrogen as the refrigerant. This explains why some air separation units with a high production of liquids, and which use large nitrogen refrigeration recycles, also achieve close to 95% recovery of argon. In the past, the point of argon withdrawal from the upper (low pressure) column was not determined with any degree of accuracy. Rough calculations combined with past experience enabled engineers to guesstimate a few outlet nozzle locations, and then hope that one was the right one. Today, though, a reasonably precise location can be precalculated. With high recovery in mind, it is important to maximize the amount of liquid nitrogen reflux in the lower (high pressure) column to strip all argon and oxygen from the rising airstream and to accumulate them in the sump of the tower. This so-called rich liquid, rich in oxygen, also contains all argon that was in the process air. It is withdrawn, subcooled through an expansion valve, and sent to the upper (low pressure) column for rectification.
4.3.4 RECOVERY PROCEDURE
AND
EQUIPMENT
It should be noted that the upper or low-pressure column of an air separation plant is never in a steady state (Figure 4.4). The prepurification system upstream of the separation equipment—reversing exchangers, molecular sieve adsorption—plays an important role in the argon recovery. Each time there is a reversal of the exchangers or a change in the airflow due to the pressurizing and depressurizing of the adsorber vessels, there is also an upset in the upper column, which in turn affects the mixture of oxygen–argon–nitrogen at the point of takeoff. This loss of argon recovery will vary with the magnitude and frequency of the upsets caused by the prepurification system.
ß 2006 by Taylor & Francis Group, LLC.
Crude argon
Crude argon column
Low pressure column Rich liquid filters
Lox Filter
Vaporizer condenser Liquid Nitrogen Lox Pumps
To and from heat exchangers
Oxygen
L o x
P u r g e
Nitrogen reflux cooler
High pressure column
Rich liquid cooler
Air Rich liquid Nitrogen
FIGURE 4.4 Crude argon recovery from standard air plant. (Courtesy of F.G. Kerry, Inc. With permission.)
From the upper (low pressure) column, a stream of argon is withdrawn and piped to a crude argon fractionation tower. The stream can contain 9% argon and 0.01% to a few vppm of nitrogen, and the remaining oxygen. To separate this large amount of oxygen, great number of trays are required; the higher the recovery of argon, the more trays are needed. The crude argon tower can thus be very tall and may require a separate cold box parallel to the main cold box. The tower is fitted with a reflux condenser developed by Georges Claude12 in the early 1920s. (Originally this piece of equipment was called a dephlegmator.) The shell side contains subcooled rich liquid. Oxygen in the argon-bearing stream is condensed and passes down to the sump whereas the slightly more volatile argon is condensed higher in the
ß 2006 by Taylor & Francis Group, LLC.
column and collected in a specially designed shelf. At this shelf, the colder liquid argon acts as liquid nitrogen in the shelf of the air plant’s high-pressure column, but in this case to strip oxygen from the rising argon-rich stream.
4.3.5 SECONDARY RECTIFICATION
AND
FINAL PURIFICATION (SEE FIGURE 4.4)
The argon-rich liquid with an approximate analysis of argon 98%, with oxygen at 1% or considerably less, the balance nitrogen, and at a temperature of around 87 K, is withdrawn from the crude argon tower and sent outside to a special external heat exchanger. In this heat exchanger, the cold argon stream releases its refrigeration in countercurrent heat exchange to an entering refined warm argon stream. The term refined here means that all the oxygen has been completely eliminated. The most volatile components, namely nitrogen and some traces of argon, are removed from the top of the crude argon column, and may be routed to the low pressure of the main plant for reliquefaction and further rectification. The refined argon contains the original amount of nitrogen, sometimes less than 1 vppm oxygen, and some residual hydrogen left over from the deoxo catalytic unit. Now cold, it is routed to a trayed pure argon column, which is similar in design to the crude column. Its reflux condenser is filled with liquid nitrogen on the shell side at a subcooled liquid temperature. The refined and refrigerated argon enters the pure argon column near the sump in a closed circuit. The closed circuit coil that receives the refrigerated argon acts as a reboiler for the pure liquid argon product in the sump. As the refined argon exits from the tower, it enters a flash drum (separator) where the condensed refined argon is expanded and degassed. This side operation serves to get rid of any excess hydrogen and nitrogen, which now exits as vapor and is returned to the crude argon system for recovery. The balance is returned to the middle of the column as feed. The vapor from this feed rises to the top, enters the tube side of the overhead reflux condenser where it comes into countercurrent exchanger with subcooled liquid nitrogen in the shell side. This action serves to reliquefy the argon. Any gaseous nitrogen and residual hydrogen escape at the top of the condenser. Liquefied argon descends through the tower and acts as a reflux to the rising gaseous mixture. The small stream of nitrogen and residual hydrogen vapor passing through the condenser is either returned to the waste nitrogen exiting from the top of the main upper (low pressure) column of the air separation plant, or is vented from the top of the pure argon column through a special pressure indicator control (PIC) valve. Liquid nitrogen at the overhead condenser is fed from the pure liquid nitrogen found at the top of the lower (high pressure) column of the main air plant. It enters the sump of the pure argon tower in a closed circuit where it is maintained as a liquid, then as it exits, it is throttled slightly for subcooling and enters the reflux condenser at the top to replenish the liquid nitrogen bath. The purified argon product in the sump is now completely free of all impurities to a level of less than a few vppm for the total sum of oxygen, nitrogen, and hydrogen. The final liquid product is withdrawn and sent to a vacuum-insulated cryogenic storage tank. Vapor from the storage tank is recycled to the pure argon column for recondensation and recovery. It is also possible to recover some part of the hydrogen stream used in the refining operation by recycling the gas stream from the flash drum and separator into the crude argon system.
4.3.6 REFINING OPERATION
AND
EQUIPMENT
The recent trend of most cryogenic designers is to produce crude argon almost completely stripped of oxygen down to 1 vppm or less (Figure 4.5). To carry out this process, one must either design a very large number of theoretical trays, which is not practical, or better still, use structured packing. Otherwise the pressure drop would prove excessive. This procedure, if successful, completely eliminates the need for warm argon refining. Nevertheless, the following description is submitted, as it may prove useful as process technology.
ß 2006 by Taylor & Francis Group, LLC.
ß 2006 by Taylor & Francis Group, LLC.
F1
T1
Crude argon Nitrogen
P1
PIC
Crude argon compressor
Deoxo argon O2 AP
Separator
LSRE
PIC
dP1
Dust filter P1
P1
AIC
T1 T1
P1
T1
Dessication bottles Flame arrestor
Deoxo unit AIC
TLS
T1 T1
P1
T1
Vent Water Refrigeration cooler unit Separator Separator
T1
Atmospheric cooler
FIGURE 4.5 Oxygen removal from crude argon. (ß Air Liquide, all rights reserved, 2006. With permission.)
The refiner consists of a group of equipment to completely eliminate oxygen from the argon stream leaving the crude argon side arm column, and operating at normal temperature levels. This equipment consists of the following: an oil-free blower, a combustion tower containing a palladium catalyst, a refrigeration machine, a water separator, a dryer filled with activated alumina, and ancillary piping, valves, and controls. This entire assembly is prepackaged and skid mounted in a module for easy erection and connection to the cryogenic equipment. The argon-rich stream is compressed in the oil-free blower and injected with hydrogen in a proportion slightly higher than the stoichiometric ratio required for the removal of oxygen completely. In the combustion tower, the oxygen is consumed completely. Hot effluent from the combustion tower is first cooled by either air or water, and then chilled by a refrigeration machine. The water so formed is removed, first in a separator and then in a dual tower-activated alumina dryer. The final product is returned to the argon system where it is cooled in countercurrent heat exchange with the outgoing stream leaving the crude argon column. In catalytic combustion for oxygen removal, there is a marked increase in temperature, 808C, or higher, for every 1% oxygen content in the treated gas. For this reason, it is prudent to recirculate the effluent around the combustion chamber in order to dilute the oxygen content and to keep the temperature down to a permissible limit. Nevertheless, the temperature of the unit is sometimes maintained at 370 K to avoid any operational upsets. The necessity and inclusion of a refiner in the recovery of argon is an expensive, albeit a necessary concept in terms of capital, energy, and maintenance. It also involves the use of hydrogen, which is costly especially in large quantities. Hydrogen may be supplied in cylinders, as liquid hydrogen in bulk form, or from cracked ammonia. This last option is not recommended, however, because it increases the amount of nitrogen content to be removed in the pure argon column. The cryogenic industry had worked for years to develop a process for the cryogenic complete removal of oxygen in the crude argon column. Finally in 1989, Linde AG built an air separation plant in Germany with an argon production, which did away with the warm argon refining system. The crude argon column was filled with structured packing, and was capable of consistently removing oxygen down to less than 1 vppm. The entire air separation plant, including the argon purification, was operated automatically from a remote control center. The argon column was so tall that its casing would have overshadowed the main air fractionation tower. The project engineer, therefore, increased the elevation of the main tower in order to achieve the same top elevation for both towers. The space below the main tower had enough room to include the liquid recirculating pumps.
4.3.7 APPLICATIONS
OF
ARGON
In the 1950s, argon was sold in cylinders. Today, it is produced in liquid tonnage quantities and bulk-delivered in insulated tank trucks or even railroad cars. It is marketed in everincreasing quantities for manifold industrial applications. In the metallurgical industries, argon is used to inert molds in pressure die casting, in the fabrication of mild steel, to protect molten metal in the continuous casting of carbon steel, and it has been stated that about 90% of stainless steels produced in the United States use this argon–oxygen decarburization (AOD) process. A major application is also in tungsten-inert gas (TIG) and metal-inert gas (MIG) shielded arc welding, especially for aluminum alloys to prevent formation of nitrides. Whereas its original introduction for filling incandescent lamps is gradually encroached upon by krypton–xenon mixtures, it is extensively used as a carrier gas in the manufacture of semiconductors and other electronic components. It also finds use in the manufacture of plate glass and in fabricating hermetically sealed multipane insulated windows.
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4.4 4.4.1
KRYPTON AND XENON GENERAL
Krypton derives its name from the Greek (meaning hidden). It has an atomic number of 36, a molar mass of 83.80 kg=kmol, no known isotopes, a vapor pressure of 731 mbar at its triple point of 116 K, a boiling point of 120.5 K, and a critical pressure of 54.82 barA. Xenon, which derives its name from the Greek (meaning stranger), has an atomic number of 54, a molar mass of 131.3 kg=kmol, no known isotopes, a vapor pressure of 816 mbar at its triple point of 161.3 K, a boiling point of 202.15 K, and a critical pressure of 58.99 barA.
4.4.2
SOURCES
OF
KRYPTON
AND
XENON
Both krypton and xenon are found in the atmosphere in extremely low concentrations. Krypton is at 1.14 106 vppm, whereas xenon is at 0.086 106 vppm. If 1 m3 of air is treated, one cm3 of krypton and one-tenth of a cubic centimeter of xenon will be recovered. Nevertheless, in 1918, Claude13 suggested that krypton be used in place of argon for incandescent lamps because it would retard the vaporization of the filament. This would make it last longer and increase the output of lumens. In 1928, Claude extracted krypton from an air separation plant. Apart from trying to recover these gases from the atmosphere, it is also possible to recover them from chemical processes, which use air, and periodically have to purge a substantial quantity of gas due to a buildup of inerts. These may also include krypton and xenon, for example, ammonia purge gas.
4.4.3
RECOVERY
OF
KRYPTON
AND
XENON
In view of the growing interest of using krypton for the lamp industry, I.G. Farben14 in 1933 decided to install a krypton recovery system in one of its three large air separation plants, each having a daily capacity of 90 t of impure oxygen, which were being constructed at Leuna for the gasification of lignite. In that first industrial adaptation, it was possible to enrich liquid oxygen in terms of 1 part krypton in 1000. It was impossible to achieve a further enrichment, however, because of the dangerous increase in hydrocarbons within the liquid oxygen sump in the MC. Following this preliminary pilot plant at Leuna, two subsequent plants were built around 1938, one unit at Boulogne Sur Mer in France treating 30,000 Nm3=h of air and another unit at Ajka in Hungary, treating 22,000 Nm3=h of air. In order to recover krypton at a time when oxygen plants were of a small capacity, the units had to be designed specifically for krypton as the main product15. It was at this time that the application of aluminum-packed recuperative exchangers, better known as reversing regenerators, were used for the removal of water, carbon dioxide, and some hydrocarbons. These reversing regenerators developed by Frankl16, opened the door to the design and fabrication of large air separation units. Unfortunately, while these units could eliminate water and carbon dioxide to a large extent, they were not successful in the complete elimination of the most dangerous hydrocarbons now in existence in the atmosphere. The recovery of a krypton and xenon mixture can be a very dangerous process as it involves the reboiling or vaporization of the more volatile component, namely oxygen, thus increasing the concentration of dangerous hydrocarbons in a bath of 100% liquid oxygen. Violent explosions have occurred in the past. In some places, the auxiliary cryogenic equipment carrying out the recovery of krypton and xenon has been separated and enclosed by a concrete barrier. In an early case, the air separation plant for the production of krypton and xenon was located at a high altitude in the European Alps in order to reduce contamination by dangerous hydrocarbons.
ß 2006 by Taylor & Francis Group, LLC.
FIGURE 4.6 Oxygen generation of 2,000 MT=d producing approximately 12 M3=d of krypton and 1 M3=d xenon. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
Before World War II, an air separation process airflow of 30,000 Nm3=h would produce approximately 200 t=d of oxygen, which was then considered huge. Today, such a plant is considered as a small-prepackaged unit. In fact, nowadays a plant capacity of 1000 t=d may be considered as routine. For this reason, krypton and xenon are now considered as by-products of air separation plants with a capacity of at least 300 t=d. One plant built in the United States in the late 1980s consisting of two units, each with a daily capacity of 1550 t of oxygen, has been producing approximately 12.5 normal cubic meters of krypton and 1 normal cubic meter of xenon on a daily basis (Figure 4.6). The process cycle for the by-product recovery of a krypton and xenon mixture is somewhat similar to the recovery of argon. In this process cycle, as is described by Isalski17, liquid oxygen product is withdrawn from the MC of the air separation unit, and after passing through a well-designed hydrocarbon adsorber, is sent to an auxiliary rectification column somewhat similar in design to a crude argon column, but with a very wide diameter in order to enhance rectification. In the latter, a portion of the liquid nitrogen is sent from the topshelf of the high-pressure column of the main plant, to the top of the reflux condenser of the rectification column. This piece of equipment serves as a noncontact condenser for krypton and xenon as well as some of the oxygen all of which have higher boiling points. The remaining liquid oxygen, now enriched with krypton and xenon, is withdrawn and sent to the sump of a third vessel where the liquid oxygen is reboiled by a stream of warm process nitrogen that is sent to the sump of the main lower (high pressure) column. The remaining liquid in the sump of this third column contains a maximum enrichment of krypton and xenon, about 1000–2000 vppm, and is ready for the final refining process (Table 4.7).
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TABLE 4.7 Boiling Points of Atmospheric Gases Element
Boiling Point (at 101.325 kPa)
Nitrogen Air Argon Oxygen Krypton Xenon
77.3 K 80.9 87.3 90.2 119.8 164.9
Source: From Pauling, L. in General Chemistry, Periodic System of Elements, W.H. Freeman & Company, San Francisco, 1958. With permission.
4.4.4
REFINING
OF
KRYPTON
AND
XENON
The quantity of recovered krypton and xenon mixture to be refined even from a very large air separation unit is not very impressive. The economics of the refining process, therefore, depend entirely on the quantity of the mixture to be separated. If the quantity is very small, refining does not involve any cryogenic distillation. After the usual means of vaporization, combustion, and adsorption to remove any residual hydrogen and hydrocarbons, the krypton and the xenon mixture can be extracted by the adsorption and desorption in a silica gel trap enclosed in a bath of liquid nitrogen. For the further separation of krypton from xenon, a small cryogenic column is necessary using liquid nitrogen at the top reflux condenser, and the conversion of air to liquid air as the reboiling medium at the bottom of the column. If, on the other hand, the quantity of the crude product is sufficiently large, a cryogenic distillation tower may be desirable. In the latter process, after the mixture goes through treatment for contaminant removal and cooling, it enters a side column with a reflux condenser at the top, which is filled with liquid air. This air is liquefied by using a high-pressure air, which is precooled in an exchanger in countercurrent heat exchanger with the vaporized air from the reflux condenser. This precooled air enters the bottom of the auxiliary column and acts as a reboiler stream to the liquefied mixture of krypton, xenon, and oxygen. As the vapor rises to the top and through the tubes of the reflux condenser krypton, xenon, and some of the oxygen are recondensed and descend to the bottom where they concentrate to a mixture containing 80% krypton and 20% xenon. This mixture, after heated to a proper temperature by countercurrent heat exchanger, is sent ultimately to a high-pressure storage tank for further separation either locally, or for transportation to another site with better facilities for the separation. Even if the latter method is chosen, it may be necessary to carry it out on the day shift, or perhaps when the total quantity is sufficient, to make it viable from an economic point of view.
4.4.5
RECOVERY
OF
RARE GASES FROM AMMONIA PURGE GAS
While several units have been built for the recovery of argon from ammonia purge gas, they were designed in a period when argon was in short supply; since then none has been built. As for the recovery of krypton and xenon from ammonia purge gas, while technically possible, it is doubtful if the units can be economically viable. The average quantity of purge gas from a 1000 t=d ammonia plant is 8558 Nm3=h. From this quantity, there is a possibility of extracting approximately 0.938 Nm3=d of krypton, assuming a 100% recovery. This is equivalent to a
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production of krypton from an air separation plant having an oxygen production of 240 t=d, which is considered as a very small plant in present-day standards.
4.4.6 APPLICATIONS
OF
KRYPTON
AND
XENON
As krypton began to be recovered, it has been replacing, or has been mixed with argon for filling incandescent lamps. Krypton lowers the rate of degradation of the filament; therefore, the lamps last longer. Furthermore, because of its heavier molar mass, a smaller quantity of gas is required. In fact, even before World War II, almost all lamps in Europe were filled with krypton. With the increasing cost of energy in North America, so-called as energy saving lamps are filled with an argon–krypton mixture. Krypton is also used in high-energy physics and lasers. Another interesting use is for stimulating plant growth due to the fact that a fluorescent lamp filled with krypton can be designed to emit light of wavelength ideally suited to stimulate photosynthesis. As xenon gives off light similar in brightness and spectrum to sunlight, it is also used as an electric flash in photography, and since the year 2000, car manufacturers have been using xenon-filled lamps for the front headlights. Airport landing lights are also filled with xenon because it gives off a bright and sharp light. It also has applications in space programs.
REFERENCES 1. Timmerhaus, K.D., and T. Flynn. 1989. Cryogenic process engineering (The International Cryogenic Monographs), pp. 25–30. 2. Gomonet, E. 1952. Les Tre`s Basses Tempe´ratures, Production et Emplois, Librairie J.B. Bailliere et Fils. Rue Hauteville: Paris, 138–139. 3. Ruhemann, M. 1945. The separation of gases (The International Series of Monographs on Physics), p. 221. 4. Ruhemann, M. 1945. The separation of gases (The International Series of Monographs on Physics), p. 222. 5. Ruhemann, M. 1945. The separation of gases (The International Series of Monographs on Physics), p. 261. 6. Timmerhaus, K. and Flynn, T. 1989. Cryogenic Systems, 2nd edn. Science and Business Media, p. 361. 7. Praxair Company Publication, 2005. 8. CryoGas International. February Issue, 2002, pp. 24–31. 9. Ruhemann, M. 1945. The separation of gases (The International Series of Monographs on Physics), p. 221. 10. Ruhemann, M. 1945. The separation of gas (The International Series of Monographs on Physics), p. 222. 11. Isalski, W.H. 1989. Separation of gases (Monograph on Cryogenics, 5), 34–35. Oxford: Oxford Science Publications. 12. Ruhemann, M. 1945. The separation of gases (The International Series of Monographs on Physics), pp. 69–70. 13. Gomonet, E. 1952. Les Tre`s Basses Tempe´ratures, Production et Emplois, Librairie J.B. Bailliere et Fils. Rue Hauteville: Paris 19:140. 14. Gomonet, E. 1952. Les Tre`s Basses Tempe´ratures, Production et Emplois, Librairie J.B. Bailliere et Fils. Rue Hauteville: Paris 19:141. 15. Gomonet, E. 1952. Les Tre`s Basses Tempe´ratures Production et Emplois, Librairie J.B. Bailliere et Fils. Rue Hauteville: Paris 19:142. 16. Gomonet, E. 1952. Les Tre`s Basses Tempe´ratures, Production et Emplois, Librairie J.B. Bailliere et Fils. Rue Hauteville: Paris 19:60–65. 17. Isalski, W.H. 1989. Separation of gases (Monograph on Cryogenics, 5), 96–100. Oxford: Oxford Science Publications.
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FURTHER READING Krepsche, R.H., B.W. Birmingham, and D.B. Mann. 1968. Technology of liquid helium, National Bureau of Standards Monograph 111. Springman, H. 1982. Methods for larger recovery to meet increased demand for the argon market. AIChE Sym Ser 79 (224):12. Springman, H. 1985. Cryogenics—Principles and applications. Chem Eng 92 (9):59.
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5
Front-End Purification Systems
5.1 HISTORICAL BACKGROUND From the beginning of the industrial application of air separation by low temperatures, it was obvious that contaminants such as water and carbon dioxide had to be eliminated. Later on, it became painfully evident that from a safety point of view other dangerous hydrocarbons, especially acetylene, also had to be removed. Present day list of dangerous contaminants that may be found in the atmosphere of a highly industrial area is given in Table 5.1. The concentrations indicated, however, are to be considered as a yearly average and not as maximum. The project engineer should make a thorough study of the local atmospheric conditions. The dangerous contaminants that may be found in atmospheric air used for the generation of oxygen include acetylene, which is the most dangerous, ethane, ethylene, propane, propylene, hydrocarbons in general, oxides of nitrogen especially nitrous oxide, and even acetone.
5.1.1 PROCESSES AND MATERIALS USED
IN
FRONT-END PURIFICATION SYSTEMS
Before any discussion of processes and materials that are employed in prepurification systems, it may be prudent to define some terms that are commonly used in the industry. 1. Absorption is the assimilation of molecules into a solid or liquid substance, with the formation of a solution or a new compound. 2. Adsorption is the adhesion of molecules of a gas, liquid, dissolved substance, or of particles to the surface of a solid. 3. Adsorbent is the material, which carries out adsorption. 4. Adsorbate is the gas, liquid, dissolved substance, or particles that become adsorbed by the adsorbent. 5. Regeneration of an adsorbent involves the action (generally by heat or release of pressure or both) taken at the plant site in order to recover the adsorbent’s capacity to adsorb. 6. Reactivation of an adsorbent involves the action required by the original manufacturer of the adsorbent to recover the material’s usefulness for adsorption and regeneration.
5.1.2 ORIGINAL PREPURIFICATION The original process and material used in air separation was absorption with sodium hydroxide (NaOH), which is a white hygroscopic (water attracting) solid that readily attracts water and becomes deliquescent. It was used for the preliminary drying of air. In a 15% solution, it was used for the removal of carbon dioxide. The process involved is two towers and is in series. The first tower contained a 15% solution of NaOH for carbon dioxide and oil removal, and the second tower contained solid NaOH for drying process air. When the lump
ß 2006 by Taylor & Francis Group, LLC.
TABLE 5.1 Assuming a Site Elevation at Sea Level at a Pressure of 101.325 kPa, the Contaminants Contained in the Industrial Quality of Air Expected in vppm Continuous Average vppm Hydrogen Carbon monoxide Carbon dioxide Methane Acetylene Ethane Ethylene Propylene Propane Butane and heavier Sulfur dioxide Hydrogen sulfide Mercaptans Ammonia Oxides of nitrogen (NO–NO2) Particulate matter
10 1 400 10 1 0.2–5.5 0.1 0.2 0–0.1 0.1 0.1 0.05 0.1 1.0 0.1 2.5 mg=m3
As mentioned, the above analysis is generally found in a highly industrial area. For a more accurate analysis, the project engineer must make his own study. Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
caustic became completely deliquescent (saturated with water), and rendered inadequate for removing anymore water it was taken out and the tower is refilled with new lump caustic. The removed solution was brought up to a 15% level and was used in the first tower for the removal of carbon dioxide. Needless to say again the process was primitive, inefficient, and messy. Yet, it withstood the test of time up to the early 1950s.
5.1.3 5.1.3.1
ADSORBENTS General
Because the use of adsorbents is so closely allied with the technology of air separation, it will be useful to review some of the adsorbents in general use. These adsorbents have been in use since 1950 in various forms and temperatures for the removal of such contaminants as water, carbon dioxide, acetylene, and many other dangerous hydrocarbons, both saturated and unsaturated. Commercial adsorbents, which exhibit ultraporosity, have been used for the selective separation of gases, and included activated carbons, charcoal, activated clays, silica gel, activated alumina, and crystalline aluminosilicate zeolites. They must be evaluated with care, however, before any selection is made. With the exception of the last, namely the zeolites, the other materials do not possess an ordered crystal structure. Consequently their ˚ ). On the other hand, pores may range anywhere from 20 to several thousand angstroms (A commercial synthetic zeolites, or molecular sieves, have pores of very uniform size from 3 to ˚ , which are uniquely determined by the unit structure of the crystal2. 10 A
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5.1.4 INTRODUCTION
OF
ACTIVATED ALUMINA
With the development of activated alumina in the early 1950s, there was a sharp improvement in efficiency. The system still used a 15% caustic solution for the removal of carbon dioxide, but was then followed by a dual tower system of activated alumina for desiccation purposes. The dual-activated tower unit generally operated on an 8 h time cycle. One tower was in adsorption, whereas the other tower was regenerated with heat. 5.1.4.1 Activated Alumina3 Activated alumina is nothing more than a pure form of aluminum oxide, completely free of water. The product is supplied in small spheres in 3–6 mm in diameter and is an inexpensive as well as an efficient material for drying gases, especially process air. It can be regenerated at a fairly low temperature to elute water, assuming there is enough sweep gas. The regeneration gas must supply a driving force to remove water condensed in the pores and adsorbed on the high surface area of the activated alumina. Whereas the heat of condensation of water is not a function of the desiccant, the heat of wetting is. Activated alumina has a lower heat of wetting than molecular sieve. It therefore has an advantage in requiring less energy for regeneration. For this reason a dual bed including both alumina and molecular sieve is more advantageous in terms of energy saving and capital cost for the combined removal of water and carbon dioxide from process air. As termed by ALCOA, the dynamic desiccant capacity (DDC) of activated alumina is based partly on experience. A rule of thumb is that in a clean and dynamic drying system, it equals one-third of the static capacity. As the relative humidity of the stream decreases, it passes through the entire bed and is not used to maximum static capacity before it is regenerated. According to ALCOA, the following DDC maxima in weight percent may be used as guidelines (conservative design uses about 3=4 maximum DDC but this may be further reduced because of severe conditions)3 (Table 5.2). 5.1.4.2 Regeneration of Activated Alumina Retained liquid water may cause breakage or fracturing of the activated alumina spheres, especially in the regenerated condition rather than the saturated condition. A gradual decline in adsorption capacity is typical in process air dryers, but this condition plateaus after a while. The decline is caused by a loss of surface area and micropore volume by repeated wetting and heating of the adsorbent. Given good conditions for air and inert sweep gases, the operating life of the activated alumina should be in the neighborhood of 10 years. This estimate, however, is reduced by half if saturated hydrocarbons are present, a condition more likely to occur in the operation of air separation plants in highly industrial areas. The total quantity of heat required for the regeneration of activated alumina can be calculated and is given in Table 5.3.
TABLE 5.2 Following Dynamic Drying Capacity of Activated Alumina May Be Used as a Guidance Drying Applications Inert gas and air Saturated hydrocarbons Unsaturated
H-156
F-200
F-1
15 11 8
12 8 5
7 6 3
Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
ß 2006 by Taylor & Francis Group, LLC.
TABLE 5.3 Distribution of Heat in the Regeneration of Activated Alumina a) Heat to alumina desiccant Ha ¼ ma (Tr þ Ta ) Specific heat (Cpa ) b) Heat for water desorption Sensible heat of liquid water þ heat of vaporization þ sensible heat of water vapor þ heat of wetting Hd ¼ mw (Tv Td )Cpw þ mw DHv þ mw (Tr Tv )Cps þ DHw c) Heat to vessel Hv ¼ mv (Tr Td )Cpv d) Heat to bed supports, screens, if present Ht ¼ mt (Tr Td )Cpt e) Heat to hydrocarbons, if present Hh ¼ 1860 kJ=kg of hydrocarbons f) Total quantity of heat estimated with 70% safety factor H total ¼ (Ht þ Hd þ Hv þ Ht þ Hh)1.7 wherein ma, mass of alumina (kg); mw, mass of water (kg); mv, mass of vessel (kg); mt, mass of bed support (kg); Tr, temperature of regeneration gas (K); Td, temperature of inlet drying stream (K); Tv, temperature of vaporization for water desorbing from activated alumina, assume 395 K; Cpa, heat capacity of activated alumina 91.7 kJ=kg=K; Cpw, heat capacity of liquid water 4.16 kJ=kg=K; Cps, heat capacity of water vapor 2.08 kJ=kg=K; Cpv, heat capacity of vessel 0.41 kJ=kg=K (carbon steel); Cpt, heat capacity of bed support 0.41 kJ=kg=K (if ALCOA T-162 is used) Hw, heat of wetting maximum F-1, 32.5 Btu=lb or 76 kJ=kg of alumina H-151, 44.0 Btu=lb or 102.1 kJ=kg of alumina F-200, 43.0 Btu=lb or 100 kJ=kg of alumina Source: Courtesy of ALCOA Bulletin, 2000. With permission.
5.1.5
ZEOLITES (MOLECULAR SIEVES)4
The study of zeolite crystals dates back to 1840 when Damour reversibly dehydrated them without any change in transparency or external form. Grandjean later found that dehydrated crystals of zeolite could reversibly adsorb inorganic vapors such as iodine, mercury, and ammonia. In 1925, Weirgel and Steinhoff found that the zeolite chabazite could adsorb water vapor, methyl and ethyl alcohol but could exclude acetone and benzene. In the early 1930s, x-ray diffraction studies revealed the zeolites as crystalline materials having within each crystal a system of precisely arrayed cavities and pores. In 1932, McBain was the first to use the term molecular sieve to explain the significance of these. Molecular sieves retain adsorbates by a strong physical force somewhat similar to magnetism. In other words, when the adsorbed molecule is removed by either heat or another material the remaining crystal of the sieve is in the same chemical state as before. Desorption of the adsorbate from the sieve produces no hysteresis effect. The adsorption and desorption phases are completely reversible, and with a few exceptions the respective isothermal curves coincide completely. 5.1.5.1
Chemical Formula5
Synthetic zeolites are better than natural types because they are produced with closely controlled chemical composition and a pore size suitable for the molecules to be passed (or excluded) in the specific purification involved. Commercially produced zeolites have the following basic formula of the unit cell:
ß 2006 by Taylor & Francis Group, LLC.
TABLE 5.4 Chemical Formulae of Various Zeolites Used in Industry Type 4A Type 5A Type 13X
Na12[(AlO2)12(SiO2)12]27H2O Ca4.5Na3[(AlO2)12(SiO2)12]H2O Na86[(AlO2)86(SiO2)106]276H2O
˚ Pore size 4 A ˚ Pore size 5 A ˚ Pore size 10 A
Source: Courtesy of UOP Molecular Sieve Adsorbents, 2005. With permission.
M2=n O Al2 O3 xSiO2 yH2 O where M is a cation of n valence. Among the types that have been found to be most commercially useful are given in Table 5.4 with their unit cell formula. In all cases, the water of hydration is removed by heating before the product is ready for use. The commercial material may also include about 20% inert clay as a binder. Molecular sieves separate molecules not only by size and configuration, but they also adsorb preferentially based on polarity or degree of unsaturation. The less volatile, the more polar, or the more unsaturate a molecule, the more tightly it is held within the crystal of the molecular sieve. For example, molecules of water and methanol have positive and negative electrical poles. These polar molecules are strongly attracted to molecular sieves. On the other hand, nonpolar molecules such as methane or ethane are more weakly attracted. When a mixture of water and methane is passed over a molecular sieve, the water is readily adsorbed, whereas methane is excluded, even though both are small enough to pass through the pores easily. Water can be removed from a strongly attracted fluid such as methanol by using a molecular sieve with pore openings smaller than the methanol molecule but larger than the water molecule. When molecular sieves remove a water molecule from a process stream the reaction is exothermic. The increase in temperature depends on the affinity of the absorbate for the molecular sieve. Conversely, when water is desorbed from a molecular sieve, the reaction is endothermic, resulting in a drop in temperature, which is a hindrance to desorption. A molecular sieve can also be regenerated by a reduction in pressure. The lower the pressure, the better the regeneration. In fact, in some cases, a vacuum is employed. 5.1.5.2 Types of Molecular Sieves ˚ in diameter. At usual The Type 4A molecular sieve has a free aperture size of 4 A operating temperature, this size allows the passage of molecules with the same effective diameter. Moreover, the sodium ion can be exchanged to form other useful products. For example, replacement of sodium ions in Type A with calcium ions will produce Type 5A with ˚ . This latter sieve is highly recommended to remove dangerous a free aperture of 5 A hydrocarbons such as propane from the liquid oxygen (LOX) in the main condenser if it has been entrained beyond the prepurification system at the front end. Type 13X, which is the most widely used molecular sieve in air separation technology, has a nominal pore ˚ . It adsorbs all material with a smaller pore diameter and excludes those with diameter of 10 A a larger pore diameter. The rate at which a given material will be adsorbed on molecular sieve pellets in any operation is dependent on four variables: 1. The rate at which the adsorbates can diffuse to the activated crystals within the pellets, which means that an efficient distribution system is of the utmost importance
ß 2006 by Taylor & Francis Group, LLC.
2. The relative size of the molecules in the vapor to be adsorbed, and of the pores in the molecular sieve 3. The strength of the adsorptive forces between the molecular sieve and the adsorbate 4. Temperature The life cycle of the adsorbent will depend on the continuous retention of water and heavy hydrocarbons, such as benzene. With the proper regeneration of the sieve freeing it from the adsorbates at each cycle, the life of the molecular sieve may well last up to 15 years. If water is allowed to accumulate, however, fracturing may ensue, and the sieve may last only 5 to 10 years. The same is true if hydrocarbons are permitted to coke the sieve. The life cycle of a molecular sieve, especially that of Type 13X, can be extended substantially by the use of a layer of activated alumina upstream of the sieve. Activated alumina has a very high affinity for water and for very heavy hydrocarbons. This procedure will shorten the life cycle of the alumina but will extend that of the molecular sieve, which has a higher capital cost.
5.1.6 SILICA GEL Silica gel is a colloid of silica and resembles very close in nature to the zeolite minerals. It has a structure, which contains corridors and chambers within its aluminosilicate framework, as well as alkali and alkaline earth ions. When a mineral such as chabazite CaAl2Si4O126H2O is heated, water molecules are driven off, but the crystal structure does not collapse. Therefore, the dehydrated mineral has a strong affinity for water, and for molecules of other vapors such as acetylene. Silica gel has long been in use as a drying agent for air conditioners, packing, etc. In 1950, it was also found to be very efficient as a filter trap in the removal of acetylene from liquid air and LOX at cryogenic temperatures. Tests on various drying agents are given in Table 5.5. As noted, using activated alumina as a base reference, the ordinary silica gel is slightly superior by 7% on an equal volume basis to the manufactured Mobil Sovabead, and very superior by 50% to Sovabead on an equal weight basis. The manufactured Sovabead has certain advantages, however, in that it can be handled easily, and does not dust, which is much appreciated by the operating departments. Results of thermal shock tests, that is, changing the temperature between 383 and 78 K in 10 cycles, indicated that there was very little loss of both, the Sovabead and the standard silica gel. The loss was around 0.003% for both. Results of mechanical vibration tests, however, indicated that the Sovabead was slightly superior to silica gel, and both were greatly superior to activated alumina (Table 5.6).
TABLE 5.5 Comparison of Various Drying Agents
On an equal weight basis Adsorption in LOX Desorption at 373–388 K On an equal volume basis Adsorption in LOX Desorption at 373–388 K
Alumina (%)
Sovabead (%)
Silica Gel (%)
100
103
154
100
117
125
Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
ß 2006 by Taylor & Francis Group, LLC.
TABLE 5.6 Screen Sizes of Typical Drying Agents
Retained on 50 mesh screen At warm temperature At cold temp.
Alumina (%)
Sovabead (%)
Silica Gel (%)
91.38 99.831
99.924 99.967
99.68 99.96
Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
5.2 DESIGN OF CURRENT FRONT-END PURIFICATION SYSTEMS 5.2.1 GENERAL BACKGROUND6,7 For more than the past decade, the use of a front-end prepurification (FEP) system with adsorbents has played a dominant role in the design and construction of oxygen plants, both large and small. In fact, this option has pushed the previous use of reversing heat exchangers (Revex) into the background, the latter being mostly limited to small units less than 200 t=d, and to nitrogen-generating units. This is the case even though the present Revex cores are fabricated under vacuum and are more reliable. Overall, the FEP system has many advantages: it is more stable in operation; it is safer in providing a more thorough elimination of hazardous hydrocarbons; and it facilitates a faster turn around in periods of total defrosting, or for maintenance. On the other hand, one must accept a somewhat higher investment cost, and a marginally higher energy consumption. In contrast, Revex systems have not been able to recover as much of the air feed as pure products (only about 45%), but require 5%–10% less energy. During the 1950s and 1960s, Revex had pushed FEP systems into the background, especially for chemical and metallurgical plants that had limited use for pure nitrogen. In the early 1970s, however, a change of scale occurred in petrochemical plants: the amount of oxygen required increased substantially to a magnitude of 1000–2000 t and more per day. The increasing size of air plants required an arithmetic increase in the number of unit aluminum cores needed for Revex systems, and their application became expensive. Even more significantly, however, field performance of the large cores suffered. The old salt-bath brazing process, coupled with difficulties in maintaining quality control during the cores’ manufacture, made their use for large air plants suspect even though vacuum brazing of the aluminum cores had started to overcome this problem. Throughout this period, however, designers of prepurification units had exerted themselves to reduce the differential in capital costs and energy requirements. They had lowered pressure drop across the adsorbent bed, reduced regeneration temperature of the absorbents with the addition of activated alumina upstream, and eliminated the mechanical chiller whenever a large quantity of waste nitrogen or other inert gas was available. With the use of an evaporative water chiller (EWC) operating in conjunction with a direct contact aftercooler (DCA), it was possible to chill process air down to 283 K, which is the optimum. However, this required the use of about 50%–60% of the total process air as waste nitrogen. With regard to adsorbent vessels, pressure drop has been lowered to less than 6.9 kPa, by minimizing the bed depth and increasing the cross-sectional area in order to maintain bed volume. There is no easy solution to this problem, however, which has resulted in a variety of design configurations. It was also discovered that by using 15%–25% of the total air feed it was possible to regenerate the adsorbent bed at a reasonably lower temperature. Where high peak load demand was a problem using an electric heater for regeneration, it was also possible to use a heat accumulator to eliminate such peaks, or use the heat capacity of the processed air that is compressed (see Section 5.3.5).
ß 2006 by Taylor & Francis Group, LLC.
5.2.2
EQUIPMENT USED
The equipment currently involved in the design of FEP systems may be summarized as follows. 5.2.2.1
Precooling Units Upstream of Adsorption
For small plants, 300 t=d of oxygen and under, a standard shell and tube aftercooler downstream of the main air compressor may prove more cost effective, depending on local conditions. 5.2.2.2
Direct Contact Aftercooler (Figure 5.1)
The design of the precooling unit may be modified depending on local conditions such as air temperature, relative humidity, feed air capacity, and cooling water temperature. For example, a noncontact heat exchanger aftercooler may be utilized, thereby eliminating the lower half of the DCA tower. On the other hand, if the availability of waste nitrogen is too low, then the use of a mechanical refrigeration unit may be unavoidable. Another point to consider at the outlet of the DCA unit is a well-designed and reliable mist eliminator in order to avoid serious entrainment of water into the adsorbent beds. The material of this mist eliminator should be stainless steel or plastic, preferably with overlapping layers. If horizontal sections are supplied they should be tightly fastened to avoid any open spacing between the sections to avoid any water entrainment. As shown in Figure 5.1, a DCA consists of two sections in a single vertical column. The lower (high pressure) section receives process air directly from the last stage of the main air compressor at a high temperature. It may be trayed or packed with rings. Polypropylene should not be used because of the high temperatures encountered at the bottom of the tower. Molecular sieve adsorber Process air 100,000 Nm3/h Temperature 8⬚C saturated CW—43 m3/h Temperature—6⬚C Waste N2—57,000 Nm3/h Temperature—30⬚C
CW—42.3 m3/h Temperature —11⬚C
LIC
Waste N2—57,000 Nm3/h
1.63 m3/h
Temperature—5⬚C dry
F1
Cooling water Process air Flow—100,000 Nm3/h Temperature—87.6⬚C
LIC
CW—111.5 m3/hr Temperature —51.26⬚C 6⬚C
5⬚C
CW—111.5 m3/h Temperature—51.26⬚C
C W—110.7 m3/hr Temperature—32⬚C DCA
LIC
Cooling water EWC
FIGURE 5.1 Typical cooling system for a front-end purification system. (Courtesy of F.G. Kerry, Inc. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
Field reports have indicated that such rings have fused when exposed to the heat. Corrosion resistance is important in this section. All internals should be stainless steel. The internal walls of the tower should also be coated to prevent or retard corrosion. Two separate liquid level indicators must be used, each with an alarm, and each with an independent connection with the tower. An abnormal rise in water level would permit water entrainment into the adsorbent vessel, thereby damaging the adsorbent. (Entrainment can be permanently damaging to the adsorbent if the cooling water is treated with a too high concentration of chemical additives.) The top section of the DCA tower is mechanically similar to the bottom section, but uses a closed water circuit in series with an EWC. Depending on makeup, cooling water enters at the top at about 279 K. The process air leaves the tower at approximately 281 K. If a mechanical chiller is used, any temperature lower than 299 K is a waste of energy. If the temperature is much over 283 K a larger quantity of molecular sieve and activated alumina will be needed to prevent a CO2, or even worse, a propane breakthrough. If conditions permit, one should use an open circuit in the upper section of the DCA. As shown in Figure 5.1, two cooling water injection pumps are required, one in operation, and the other a 100% hot-wired standby spare, with all required instrumentation and controls. Two independent pressure drop indicators should be used to further protect against any entrainment of water. The upper section of the DCA tower is piped in series with the EWC. 5.2.2.3 Evaporative Water Chiller The EWC is nothing more than an open vessel packed with rings, trays, grid, or other suitable packing. Dry waste nitrogen from the cold box enters at the bottom. This nitrogen can be a major portion of the air treated in the cold box. For instance, an oxygen plant processing 100,000 Nm3=h of air will require 50,000–60,000 Nm3=h of waste nitrogen for the evaporative chiller. Recirculated water from the DCA tower or from the water-cooling tower flows from the top countercurrent to dry, cool waste nitrogen from the cold box. Alternatively, if makeup water is at a low enough temperature, an open water circuit may be used. Care must be taken, however, to avoid reaching too low a temperature in the EWC, otherwise the recirculating pump and even the packing may freeze. This has already happened in the field (Figure 5.2). The design should include stainless sieve trays or adequate polypropylene ring packing. It should also include a storage section at the bottom of the vessel to serve as a surge tank for the chilled water system. The capacity of this storage should be enough to hold at least 3 min of water storage at the maximum chilled water flow. Two chilled water-circulating pumps should be supplied; one in operation, and the other piped and installed hot-wired as a complete 100% spare. The usual shutdown alarms should be considered for high and low levels. As indicated in Section 5.2.2.3, two independent water level indicators with separate attachments to the tower are recommended. 5.2.2.4 Mechanical Refrigeration Unit (Figure 5.3) If it is decided to use a mechanical refrigeration unit, its selection should be based on high reliability and proven field service. A centrifugal or a screw-type compressor is preferable. The refrigeration fluid shall also conform to EPA standards such as HCF 134a (1,1,1,2tetrafluoroethane) or equal.
5.2.3 ADSORBER UNIT6 During the past decade, there has evolved a variety of configuration designs, but aside from the standard two-vertical vessel design, the basic few options that have proven viable in the field are worth a brief mention.
ß 2006 by Taylor & Francis Group, LLC.
AC/DC system (Airplant flow = 12,000 lb mol in.) 75 72.5 70 67.5 65
Water temparature (F)
62.5 60 57.5 55 52.5 50 47.5 45 42.5 40 37.5 1000 (b)
2000
3000 4000 5000 6000 Waste N2 flow (lb mol/h)
7000
8000
FIGURE 5.2 (a) Closed air-cooling water system. (ß Air Liquide, all rights reserved, 2006. With permission.) (b) Ideal cooling water temperature using waste nitrogen. (Courtesy of Bernstein, J., Cryogenic Consulting Services, Inc., 2006. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
Energy (kWh)
Conditions—air compressed to 6.5 bar Cooling water at 20⬚C
(C) (B)
(A) 30⬚C
25⬚
20⬚
15⬚
10⬚
5⬚
0⬚
(A) Shows the increase in energy for every degree in cooling of compressed air, using external refrigeration. (B) Shows the drop in energy requirements by using a DCA (C) Shows the new saving of energy between the two systems. After the compressed air reaches around 8⬚C there is no further saving of interest
FIGURE 5.3 Comparison of air cooling using mechanical refrigeration and a direct contact aftercooler. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
5.2.3.1 Standard Design The so-called standard design with dual vertical vessels is still used for oxygen plants up to a capacity of 300 t=d, or to a process airflow of 45,000–50,000 Nm3=h, assuming that the air pressure is at approximately 5–6 bar absolute, and that feed temperature is around 283 K. This basic design may still be used for larger capacities of oxygen production if the process air pressure is increased, for example, in the application of oxygen in integrated gas generation plants. In the latter case, however, one must take into account the critical velocity of the process air flowing through the adsorbent beds. A number of manufacturers still turn out prepackaged units of the basic design as illustrated in Figure 5.4a, which operate satisfactorily in fairly large capacities. Two vertical vessels are mounted on a skid module that contains all necessary switch valves and timers. These vessels are alternately on- and offstream: the off-stream was used to elute water, CO2, and various hydrocarbons including acetylene that have been removed from the process airstream during the adsorption phase, whereas the on-stream vessel is on adsorption. 5.2.3.2 Multiple Vertical Vessels As shown in Figure 5.4b, it has been possible to add two extra vessels operating in conjunction with a single set of switch valves. With this design, it has been possible to increase the capacity of an existing air plant from 300 to 600 t=d with no operating problems whatsoever.
ß 2006 by Taylor & Francis Group, LLC.
Waste nitrogen vent
PIC
Waste nitrogen from cold box Reactivation heater
T1
Process air to cold box
Vent P1
(b)
T1
Process air from cooling system
FIGURE 5.4 (a) Standard twin vertical vessels with horizontal adsorbers for a 300 t=d oxygen plant. (ß Air Liquide, all rights reserved, 2006. With permission.) (b) Basic vertical-vessel adsorber design. (Courtesy of F.G. Kerry, Inc., 2006. With permission.) continued
ß 2006 by Taylor & Francis Group, LLC.
Waste nitrogen vent
Waste nitrogen from cold box Reactivation heater Process air to cold box TI
Vent
(c)
Process air from cooling system
FIGURE 5.4 (continued) (c) Basic adsorber unit of vertical vessels with possible option for multiple vessels. (Courtesy of F.G. Kerry, Inc. With permission.)
Another variation of this four-vertical vessel design as shown in Figure 5.4c is to have the four vessels operating in a sequence. This design has been a successful option for many years, and has been used for units of airflows up to 125,000 Nm3=h operating at 7 bar, that is, a production of 700–800 t=d. Whereas this option involves a higher number of switch valves, the valves are of a much smaller diameter, resulting in a lower total purchase cost. The vertical vessels, mainly of a smaller diameter, also have a thinner wall thickness, therefore cheaper to roll and to fabricate. Moreover, equipment such as vessels, piping, switch valves, and timers, can be factory assembled into one prepackaged skid-mounted module complete with instrumentation, tested, and shipped to the site as a single unit. It also has an operating advantage because although it has more operating bumps in the upper (low pressure) distillation column than the two-vertical vessel adsorber design (due to the operation of the multiple switch valves), the bumps are much less severe, and the fractionation is therefore somewhat smoother. This is apparent in the resultant higher recovery of argon. Another option that has been used successfully in the field is a three-vertical vessel design, also operating sequentially. It has been applied in a plant producing 1500 t=d of oxygen. 5.2.3.3 Horizontal Vessels (Figure 5.5a) Several designers have opted for this configuration with satisfactory results, at least for very large plants (over 1000 t=d). A horizontal vessel orientation, however, increases the mechanical complexity of the grid platform supporting the adsorbents. Not only must the grid support the dead weight of the molecular sieve, but it also must withstand the cyclic thermal stresses resulting from temperature excursions from 423 to 523 K, down to 283 K, and
ß 2006 by Taylor & Francis Group, LLC.
LBOUT
LEGRI
LGFIL
A
B
DRPIQ DRFIL
Air NR
DIBOU
B Steve
Molecular
Activated alumina
LGGRI
Air
NR
r su
Filtra t
A
ion
Cut B.B
(a)
FIGURE 5.5 (a) Molecular sieve adsorber horizontal vessel and horizontal beds. (b) Torn upper screen filter due to excessive fluidization. continued
ß 2006 by Taylor & Francis Group, LLC.
FIGURE 5.5 (continued) (c) Disturbed surface of molsieve bed due to excessive fluidization. (Courtesy of F.G. Kerry, Inc. With permission.)
back again. In addition, the total assembly has to survive dynamic stresses caused by depressurization and repressurization from 6–7 bar to atmospheric pressure, and back again. Obviously, various members making up the grid will move by thermal expansion and contraction, but the movement must be such that small granules or beads of adsorbent will not fall through the grid or be ground to dust. The quantity falling through the grid must be kept extremely small—no more than a few grams per day. Another problem, but by no means less difficult, is how to distribute the process air uniformly along the large cross-sectional area of the bed. The solution chosen originally by almost all engineers was to let the process air enter from a bottom nozzle at full pressure, assuming that one simple wide baffle plate would distribute it equally across the area of the bed. They ignored a key fact, as memorably stated by an experienced designer, Layton Kitterman, ‘‘Fluids flow according to the laws of physics, and not according to the wishes of the designer.’’ Only in a few cases uniform flow has been obtained at a velocity that did not fluidize the bed, and then only when adjustable guide vanes or multiport inlets were employed. These solutions were only moderately successful, and also added significantly to the capital cost. Figure 5.5b and Figure 5.5c show results of fluidization. 5.2.3.4 Radial or Concentric Design7 Still another option that has been introduced successfully during the past decade involves two vertical vessels each with two concentric adsorbent beds (Figure 5.6). Each adsorbent bed is held together between two vertical cylindrical screens suspended from the top of the tower. Process air enters at the bottom and is directed toward the annular space at the wall of the vessel. It then travels horizontally from the wall to the center of the vessel. The process air collects in the central core of the vessel, and exits at the top. Waste nitrogen for regeneration enters centrally at the top of the tower, travels horizontally first through the molsieve, then through the activated alumina, and exits at the bottom. The advantages are obvious. The screens expand and contract freely during thermal cycles. The adsorbents can be installed in a thin bed, and tower height can be made as tall as necessary for the required
ß 2006 by Taylor & Francis Group, LLC.
Air
N2
n Radial flow double bed Alumina Molecular sieve
N2
Air
FIGURE 5.6 Vertical radial flow vessel. (ß Air Liquide, all rights reserved, 2006. With permission.)
capacity. Only simple bottom and top baffles may be needed to make air and gas flow uniformly. No guide vanes are required. Yet all is not ideal. Because the design imposes some rigid and uncompromising procedures for fabrication, labor costs are high. Therefore, the concept may be expensive for small plants. Although units have indeed been built for plants as small as 600 t=d of oxygen, for those below 800 t=d, economics may become questionable. Each case has to be studied individually. In summary, one may say that up to 700 t=d, the market belongs to either the two-vertical vessel design or the prepackaged four-vertical vessel configuration, the latter operating in sequence. For oxygen units with a capacity of 700 up to 1500 t=d, the designer may choose between the three-vertical vessel option, the horizontal vessel configuration, or preferably, the radial design. Beyond 1500 t=d, the radial design has a very definite advantage in all respects.
5.3 PROCESS OPERATION 5.3.1 ISOLATION VALVE DOWNSTREAM
OF
FEP (FIGURE 5.7)
It is of the utmost importance to have an isolation valve between the adsorbent system and the cold box to make sure that the main process airflow is introduced slowly into the equipment, especially after a temporary shutdown when cryogenic liquids are still in the distillation columns. During initial start-up it would be foolhardy to subject the equipment
ß 2006 by Taylor & Francis Group, LLC.
ß 2006 by Taylor & Francis Group, LLC.
HIC 225
T1 240
ZIC 225
FIG 427
HV 225
HIC 223
From adsorber vessels
Legend DCS shared display in control room Discrete instrument in control room Shutdown signal
ZIC 223
1"
HV 223
TI
Temperature indicator
FIC
Flow indicator control
ZIC
Position indicator control
HIC
Hand indicator control
Modulating valve
HV
Hand valve
Check valve
Note:
Main process control valve does not open until bypass valve allows pressure down stream to build up to a permissible level
Permissive signal Solenoid
Venturi flow meter Piping insulation
FIGURE 5.7 Control system at discharge of adsorber vessels. (ß Air Liquide, all rights reserved, 2006. With permission.)
To cold box
to the full pressure and flow of process air instantaneously. Secondly, after a brief shutdown due to a compressor stoppage, the liquids on the trays in the distillation columns will dump to the bottom and create a large cold sink. If in this situation a large flow of air is permitted to enter the lower column, it will condense and create a partial vacuum. This in turn will increase the airflow velocity, which will create fluidization in the adsorbent beds, as well as upset the air compressors. This condition can be controlled by the use of the isolation valve with a bypass system between the FEP unit and the cold box. Alternatively, the isolation valve can be programed so that if there is an emergency shutdown, the valve will shutdown immediately, preventing cold air from backing into the warm section. During the restart period, warm airflow will ramp-up slowly in carefully controlled increments until the downstream pressure is almost equal to the design upstream pressure.
5.3.2 ADSORPTION KINETICS The adsorption kinetics of CO2, water, and hydrocarbons in process air on molecular sieves is not as simple as just letting the adsorbent in one vessel become loaded with contaminants, and then switching over to the next regenerated unit. Contaminants in the bed are in continuous movement, and the movement of each varies precisely according to the affinity of the adsorbent for it. Concentrations are continuously changing. As shown in Figure 5.8a, each component will establish a wave front and these will move as components and are displaced by more adsorbed material (Figure 5.8b and Figure 5.8c). Adsorption of hydrocarbons by molecular sieves generally increases with increasing molecular weight in a homologous series. Unsaturated compounds are more avidly adsorbed than alkanes. The affinity of a commonly used molecular sieve for various compounds of ˚ decreases in order: water, butanes, propylene, acetylene, molecular size less than 10 A ethylene, propane, CO2, ethane, and methane. Ammonia, ethanol, and mercaptans are also adsorbed more strongly than CO2. This does not imply that contaminants are totally removed. For example, this molecular sieve takes out only 89% of propane and 97% of ethylene. Methane is hardly removed at all, and exits with the process air. With the exception of the strict specification for aviation oxygen, methane poses no problem for industrial oxygen uses. The adsorption of water by molecular sieves is exothermic, and as the bed temperature increases, therefore, the adsorption of CO2 decreases. Conversely, the desorption of water is endothermic, so that more heat is required for desorption. Also note that if too many liquid water particles are entrained into the bed, adsorption will be impaired and fracturing of the beads may occur. The resulting dust may enter the primary exchanger system where it will be difficult if not impossible to remove. In one case, for an oxygen unit (1200 t=d), after 20 years of operation, the dust is still in the primary exchangers. During the adsorption phase, it is important, not only to make sure that the process air is diffused evenly over the entire surface of the bed, but also that its flow velocity through the bed is less than 20 cm=s, preferably closer to 10 cm=s. From empirical studies in the field, the minimum quantity of molecular sieve (Type 13X) is 50 kg per 1000 Nm3=h of process airflow per 1 h of adsorption. In circumstances where the atmosphere is heavily contaminated with propane it may be wise to increase the quantity of adsorbent. Unfortunately, this suggested increase is still in the realm of guesswork, but an increase of 10% is a reasonable minimum.
5.3.3
REGENERATION CONCERNS
The complexity of the complete regeneration of a molecular sieve adsorbent has not been fully appreciated by many designers. Prepurification units for smaller plants were often designed for 8 h duty, which fitted well with the 8 h labor shifts. Regeneration temperatures were
ß 2006 by Taylor & Francis Group, LLC.
Position of hydrocarbons at the beginning
Position of hydrocarbons at the end
C3H8 C2H6
C2H2 C3H6 C4 H2O
C2H2 C3H6 C4
CO2
C2H6
Concentration
Length of the adsorber bed
CH4
CH6 C2H6
C3H8 C2H4
CO2
H2O
(a)
Flow direction
Vppm C2H6 in main condenser 200 180 160 140
Without LOX adsorber With
120 100 80 60 40 20
h
0 (b)
1.0
2.0
3.0
4.0
FIGURE 5.8 (a) Position of hydrocarbons in the adsorber bed. (b) Increased C2H6 concentration in the LOX of the main condenser at a content of 10 vppm C2H6 in the process air. continued
generally around 523 K. With the advent of large air separation plants, however, and with the introduction of automated control, such designs are no longer viable. Since large oxygen plant operators do not have a need for large volumes of pure nitrogen, this offers a designer the opportunity to use waste nitrogen to supply large quantities of heat for regeneration, without using high temperatures. It is now also possible to use waste steam at 3 bar and 407 K, or even heat from the intercoolers of the main air compressors at 353–363 K. The volume of waste nitrogen required for regeneration will be 15%–25% of the total process air, depending on regeneration temperature chosen.
ß 2006 by Taylor & Francis Group, LLC.
6
5
C3H6
Molecules in each cavity (number)
C2H4 C2H2 4
C4H8⫺1
3
2
C3H8
1
C2H6 CH4
0 100 (c)
200
300
400
500
600
700
Pressure (mm)
FIGURE 5.8 (continued) (c) Adsorption of hydrocarbons by zeolites is much greater for unsaturated hydrocarbons whose molecules contain double or triple bonds. From top to bottom the curves show adsorption (at 1508C) of propylene, ethylene, acetylene and isobutylene (unsaturated), and propane, ethane, and methane (saturated). (Courtesy of Linde BOC Process Plants, 2005. With permission.)
The lower the regeneration temperature, however, the more difficult it may be to remove the adsorbed water. Residual water will build up and impair CO2 adsorption very quickly. It is prudent, therefore, to install an auxiliary heater, preferably electric even though it may be used only once a year or even less to prevent excessive accumulation of water. Since water is more strongly adsorbed than any other contaminant, it is a prime suspect when operational problems occur in prepurification systems. To minimize water’s deleterious effects, the following options are recommended: 1. Use a well-designed DCA with high contact efficiency and low entrainment. 2. Employ a good mist eliminator at the top of the aftercooler, or an efficient water separator, or both. 3. Eliminate all the water before the air feed reaches the adsorbent bed by using activated alumina or an equivalent material as a predryer immediately upstream of the molecular sieve bed. This recommendation is very important. Using an activated alumina predryer will also eliminate the harmful effects of sulfur dioxide, which can be entrained by water. Should entrainment of acids be a problem, a molecular sieve suitable for removing them should be installed upstream of the main adsorbent. Experience
ß 2006 by Taylor & Francis Group, LLC.
with dual beds has shown that the life of the main adsorbent can be extended to well over 10 years. Furthermore, as water can be desorbed from alumina with a low heat input, a lower total quantity of heat will be required for a dual-bed system than for a single-bed unit. A final and not unimportant advantage of the dual-bed system is that less molecular sieve is needed in the main bed. With regard to the amount of activated alumina needed, one has to calculate the quantity of moisture contained in the main process airflow leaving the DCA tower, assuming a state of saturation at its exit temperature and pressure. In proportion to the mass weight of the molecular sieve, the quantity of alumina may be as low as 25%, or as high as 75%. These figures are entirely empirical, as no one has developed an algorithm to equate the necessary factors.
5.3.4 WARNING
AGAINST
EXCESSIVE HEAT
DURING
REGENERATION
The heater or the heating element should be switched off before the outlet temperature of the waste nitrogen reaches its maximum. The heat front will continue working its way through the bed (pulse heating). This will also save energy. If the heater is shutdown too late, however, there will not be sufficient time to cool the bed. After late switching, the process air will continue to flow through a still hot alumina bed, and it will not adsorb water effectively. This condition will load up the molecular sieve with water and reduce the efficiency of the molecular sieve for adsorption of CO2, C2H2, and other dangerous contaminants.
5.3.5 HIGH-PRESSURE VESSEL REGENERATION Another factor for regeneration not yet fully appreciated is the effect of pressure, especially in the range of 10–20 bar. At such high process air pressures, the thick walls (25–50 mm) of the steel vessels containing the adsorbents act as a heat sink, reducing the overall heat energy of the regeneration gas. This condition gets worse as the regeneration gas reaches the alumina at the lower level, and will impair elimination of adsorbed water. As noted, if water is allowed to accumulate in the adsorbent it will harm the overall adsorption phase. To remedy this condition, it may be necessary to incorporate an insulating material (possibly up to 50 mm) within the vessel to minimize the escape of heat through the vessel walls. To eliminate this condition completely it is advantageous to use the radial design so the regeneration gas first passes through both adsorbents beds before contacting the vessel walls.
5.3.6 REGENERATION OPTIONS
FOR
FEP UNITS
5.3.6.1 General If heat is to be used for regeneration of the adsorbent beds, any source of heat can be used: electricity, natural gas or propane, high- or low-pressure steam, hot water from the compressor coolers, or even heat from the compressed process air before the final aftercooler. What is important is constant energy availability and temperature stability. With regard to supply and degree of control, experience has shown that electricity is the most reliable source. Highpressure steam, if available, requires a fairly complex system of supply and a high degree of maintenance. Waste steam at 3 bar is more than hot enough to fit the need. In 1989, Isalski8 published an equation to include the heating and cooling times for a given regeneration gas flow: the equation is as follows: u ¼ th þ tc þ tx
ß 2006 by Taylor & Francis Group, LLC.
(5:1)
and tc ¼ (WB CPB )=(GCPG ) ln{(T1 T3 )=(T2 T3 )}
(5:2)
where u is the total on-stream time for 1 bed (1=2 NEMA cycle time), th is the heating time, tc is the cooling time, tx is the changeover time (including depressurizing, equalization, pressurization), Wb is the weight of bed, CPB is the specific heat of bed, T1 is the initial bed temperature, T2 is the required bed temperature, and T3 is the coolant temperature. The above equation, however, does not take into consideration the amount of water or any other contaminants on the bed because it quantifies the cooling period after the water has been driven off. The quantity of the latter may be as high as 35%. Moreover, it does not include such factors as high peak demands if electricity is used, the cost of fuels such as propane if electricity is out of question, and the influence of the metal mass (heat sink) of the vessel itself, especially if the process air is delivered at a high pressure. If a high temperature above 425 K is specified, the use of a heat accumulator has been found useful to reduce energy costs. Because the regeneration phase requires only 25%–35% of the connected cycle time, average energy consumption should be about 35% of the connected energy rating. Therefore, the use of a heat accumulator has been especially economical when employing electricity for the heating phase. The accumulator is a vessel packed with either copper metal or quartzite stones. To avoid high peak demand penalties, designers have used heat accumulators in series with electric heaters. They have been used for air separation units with a capacity of up to 600 t=d of oxygen. This fixture has long been replaced by the use of a standard heat exchanger also called an economizer to provide a countercurrent heat exchange between the hot process air from the main air compressor and the cold waste nitrogen from the primary heat exchangers of the cryogenic unit (Figure 5.9). The other factor involved the use of higher pressures for the main process air, which increases the thickness of the vessels, if the standard design is used. This problem has been discussed in Section 5.3.4, and in these circumstances one may have to apply an insulation material between the inside surface of the vessel and the adsorbent material. Another system is to use multiple adsorbent vessels, as previously mentioned, and to operate each vessel in a well-defined sequence. Waste nitrogen being discharged from vessel A being regenerated flows through and warms up vessel B that has already been depressurized,
Preheated waste N2 to main heater
Cold waste N2 to heat economizer
Heat economizer
Warm process air from air compressor
Cooled process air to DCA lower section
FIGURE 5.9 Using heat from air compression to preheat waste nitrogen for the regeneration of adsorbers. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
FIGURE 5.10 Multivessel absorption system. (Courtesy of Chemical Design, Inc., 2005. With permission.)
and so on. Generally four vessels are used for this system, each one operating at a different cycle phase. In this manner no heat from waste nitrogen leaves the system (Figure 5.10 and Figure 5.11). Three vessels operating sequentially have also been used. A thermal pulse type
Operational sequence of FEP system at GPCG Adsorption--- 1.8 h Note: The sequence is staggered 0.52 hours per Depressure---0.1 h vessel in order to maintain a constant flow Heating--------0.52 h of heat and thus conserve energy. This factor Cooling--------0.95 h is an equal divider durring every four sequences. Depressure---0.15 h (The final timing has not yet been settled in the field.) 7/04/84 Total------------3.52 h 0:0
1:0
2:0
3:0
4:0
5:0
6:0
7:0
Hours
A H C A
DP PR
A H
C A
A H C
A
A H C A
FIGURE 5.11 Generation of multivessel absorption system. (Courtesy of Chemical Design, Inc., 2005. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
of regeneration has also been used to economize energy consumption. Regardless of the regeneration temperature selected, it is possible to adjust the heating element so that it is shut off before flowing hot gas reaches the end of the adsorbent bed. The heat wave, or pulse, will continue to move through the bed, and regeneration will be complete. An energy saving of approximately 25% may be realized with this procedure. Of course, the timing of the sequence has to be precalculated precisely. Another means of saving energy is to implant analytical probes into the adsorbent beds and monitor the CO2 already adsorbed. The regeneration phase can then be started only when the CO2 reaches a precalculated concentration, not on a time cycle. In all of those options, it is assumed that a dual adsorbent system is used with an activated alumina bed located upstream of the molsieve bed. One may conclude, therefore, that there is no single equation that can encompass all the variations to regenerate zeolites with an optimum economy.
5.3.7 SUMMARY In summary, two criteria must be satisfied in the design of successful prepurification units for regeneration purposes: 1. Regeneration gas must be diffused uniformly over the entire surface of the bed to prevent accumulation of residual water and the dusting of the sieve bed. This point demands great emphasis, as it has been a major problem for almost all designers. 2. Regeneration gas must have sufficient pressure to overcome the pressure drop of the absorbent bed, heater, piping, and valves. Otherwise, back pressure will develop in the upper, or low pressure, column of the air separation unit, thereby lowering the distillation efficiency. This problem can be overcome either by increasing process air pressure slightly, by using a small booster blower for waste nitrogen (which is generally anathema to plant operators), or by increasing the number of passages for the waste nitrogen stream in the front-end heat exchangers. In any case, there will be a slight increase in the total energy consumption or investment. 3. For large air separation plants, over 600 t=d of oxygen, consider the use of the radial design, because it has a low pressure drop.
5.3.8 OPERATIONAL TIME CYCLE (FIGURE 5.12) The total operational cycle, also known as the NEMA cycle, should cover both the adsorption and desorption phases, and should include . . . . . . . . . .
Equalizing pressure in both vessels Opening flow in the second vessel Isolating the vessel to be regenerated Depressurizing the vessel to be regenerated Heating waste nitrogen to the required temperature Passing hot waste nitrogen through the adsorbent bed Cooling the regenerated vessel to ambient temperature with cool, dry waste nitrogen Pressurizing the vessel Diverting process airflow through the regenerated vessel Using the prepared vessel for a complete adsorption cycle
All the steps are sequentially calculated, and then modified to suit field conditions. If they are programed and computerized, it is wise to install a manual override in the system. The total
ß 2006 by Taylor & Francis Group, LLC.
Both vessels in operation
Vessel A Operating
Pressure
10 Bar
1 Bar
Heat
Cool
120⬚C
30⬚C
Regeneration time 240 min
480 min Time
FIGURE 5.12 Typical regeneration heating time 185 min minimum. Cooling time 118 min and depressurization 55 min. (ß Air Liquide, all rights reserved, 2006. With permission.)
NEMA cycle time may vary between 4 and 16 h. A realistic choice is 4 h, because it reduces the quantity of adsorbent needed yet does not increase maintenance cost unduly. In the 4 h cycle, 2 h of adsorption and 2 h of regeneration take place. The final choice, of course, depends on the supplier’s design, and experience with similar plants.
5.3.9 PREPURIFICATION ADSORBENT UNITS AND OPERATING STABILITY Stability of operation of the adsorption system is of the utmost importance. Because concentration of atmospheric contaminants continually fluctuates, the system imposes a need for constant vigilance. Its design involves the continuous and complete removal of not only carbon dioxide (for which it is basically designed) and water, but also of dangerous contaminants such as acetylene, ethane, ethylene, propane, and other hydrocarbons. Removal of CO2 is monitored either by an in-line gas chromatograph or a gas cell equipped with a dispersive or nondispersive infrared analyzer. These instruments have to be recalibrated at least once a week. Instrumentation should be capable of detecting CO2 down to at least 0.01 vppm. If it is observed that a serious quantity begins to breakthrough the adsorbent, action must be taken immediately. Otherwise dangerous hydrocarbons may be entrained. In fact, in a gaseous oxygen plant with minimal production of LOX, the mandatory concentration of carbon dioxide after adsorption should not exceed 0.01 vppm. An excessive concentration in the purified airstream can lead to an accumulation of solid CO2 on the main condenser–vaporizer surface and valves, especially in a gas-producing plant. This will block uniform flow of LOX, creating pockets of dry evaporation and dangerous concentrations of residual hydrocarbons. In LOX pumping, the entrained carbon dioxide will be carried away by LOX to be vaporized in the LOX pumps. In a LOX-producing unit, the CO2 will be carried out into the storage tanks (Table 5.7). 5.3.9.1 Improving Operating Stability To repeat, it cannot be overemphasized that it is of the utmost importance to operate any adsorbent purification system with high operating stability. If any instability occurs, one can improve the systems stability as follows: (a) Small breakthrough (1–3 vppm)
ß 2006 by Taylor & Francis Group, LLC.
TABLE 5.7 Contaminant Breakthrough in Molecular Sieve Beds Methane Ethane Carbon dioxide, propane, ethylene Acetylene, propylene Butanes Water
Immediately Almost immediately Almost simultaneously, but in the reverse sequence Later Much later Very much later
Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
This situation may be caused by the following: 1. An insufficiency of adsorbents (molecular sieve or alumina) to take care of impurities such as carbon dioxide and water 2. The adsorption phase is too long, and permits the carbon dioxide to breakthrough 3. The reactivation temperature is too low, its timing is too short or too long (overheating) Action to be taken: . . .
Reduce the process airflow slightly Shorten the adsorption time and observe the results after several complete NEMA cycles Increase the reactivation temperature
(b) Sudden serious breakthrough (5–10 vppm) Action to be taken: .
. . .
.
5.4
Strictly carry out the same action as mentioned for a small breakthrough, but more pronounced Reduce the process airflow Check the overall sequence and timing of each cycle Check the LOX product in the main condenser for any sign of milkiness. If it is pronounced, the product may require dumping Take a sample of the LOX product to the chemical laboratory and analyze it for any sign of acetylene, ethylene, propane, etc.
SAFETY
The low-temperature industry has a commendable safety record, mainly due to the extreme caution engineers and operators have used in designing and operating air separation plants. When the rare explosion had occurred, it was blamed on the presence of acetylene in the feed air or on acetylene-generating equipment in the vicinity. This was disproved in 1956; however, when an explosion in a large oxygen plant took place, but no traces of acetylene was found8. Investigation showed that the explosion was due to accumulation of nitrogen oxides, ethane, ethylene, and other paraffins and olefins. This incident and similar ones impelled designers of Revex units, which at the time were favored, to employ silica gel adsorbent to remove hydrocarbons from the oxygen-rich liquid downstream of the lower (high pressure) column in the cold box, and to add an extra LOX guard filter, also filled with silica gel,
ß 2006 by Taylor & Francis Group, LLC.
to remove any remaining traces of dangerous hydrocarbons from LOX product in the main condenser. At the same time, it was recommended that a small stream of LOX (0.5%–1.0% of total product) be withdrawn from the main condenser. These design modifications reduced the number of accidents to almost zero, but they also increased investment and operating costs. With the resurgence of FEP systems, however, safety design has developed new parameters. Dangerous contaminants in the air are still present and in ever-increasing quantities. Their removal requires a change in the concepts as well as in procedures, which differ from those employed with Revex systems. Contaminants breakthrough a molecular sieve bed in reverse order of their adsorption.
5.4.1 HYDROCARBON BREAKTHROUGH9 Because acetylene, propane, propylene, and butanes breakthrough the bed more slowly than CO2, by carefully monitoring CO2 breakthrough it is possible to prevent the passage of dangerous hydrocarbons into the cold box. But the system is by no means foolproof: 3%–10% of propane and ethylene will not be adsorbed, methane breaks through immediately, and ethane soon after. Unfortunately, trying to prevent methane breakthrough is not economically feasible. Thus, residual traces of methane, ethane, propane, and ethylene will remain in the LOX located at the main condenser. Solubilities in LOX at process temperatures are about 1 mol % for propane, 2 mol % for ethylene, and 12 mol % for ethane and methane. The plant designer must also consider possible atmospheric inversions, which may take place especially in coastal areas, and which may increase pollutant concentration at the air plant intake equipment beyond design. Despite care in design and in operation, errors may allow CO2 to breakthrough in FEP systems, plugging up passages in the main condenser. Small, isolated, and quiescent pools of LOX may also form and if ambient air is heavily polluted, as is the case in refinery and petrochemical areas, these pools will develop a high concentration of ethane and propane. An enormous quantity of CO2 can accumulate in a cold box from even an insignificant trace passing through the FEP unit. For very large plants, the situation can become horrendous, as seen in Table 5.8, which is based on a plant producing 1000 t=d of oxygen with a process airflow of 150,000 Nm3=h, and an on-stream factor of 98% (8585 h). Ethane, on the other hand, has different properties; above 110 K liquid ethane is completely miscible in LOX. Between 110 and 90.18 K (the boiling point of LOX at atmospheric pressure), the two are partially miscible. Two phases can appear in the main condenser, one rich in ethane and the other rich in oxygen. Propane and propylene show behavior similar to that of ethane10.
TABLE 5.8 Annual Accumulation of CO2 Entrained in a Plant Producing 1000 mt=d of Oxygen CO2 vppm Concentration 0.10 0.25 0.50 1.00
Weight (kg=y) 253 632 1265 2530
Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
ß 2006 by Taylor & Francis Group, LLC.
Also, occlusion by a relatively soluble material such as ethane can occur in a less soluble compound that is precipitating from solution. Once precipitation has taken place, the ability of purging to remove deposits is questionable. It will depend on settling of solids in quiescent zones, adhesion to tube walls, and rates of resolubility. This means that less soluble materials can cause dangerous situations even through their concentrations are well below the recognized solubility and homogeneous flammability limits11,12.
5.4.2 SAFETY ADD-ONS The addition of an auxiliary (pressure) vaporizer can contribute to the safety of an FEP system (Figure 5.12). It will also cut down energy demands of the oxygen compressor. Conditions in this auxiliary vaporizer are more favorable for volatilization because of slightly higher vapor pressures and slightly greater turbulence. It is necessary to keep in mind, however, that hydrocarbons have a tendency to accumulate in the remaining LOX at the bottom of the auxiliary vaporizer. This location should be monitored very closely. The installation of a LOX guard filter is also strongly recommended. Many engineering firms do not offer this additional safety feature, believing that it is unnecessary. Its cost, however, is insignificant compared with the total investment of an air plant, so it may be considered as an inexpensive insurance policy. It should use a calcium silicate molecular sieve, similar to molsieve Type 5A, and not the one of silica gel. The latter substance is efficient only for removing acetylene, but has very low adsorption for ethane, particularly in the presence of CO2. Silica gel is only partially effective for C3–C5’s compounds10. For an unusual quantity of propane the only rule of thumb used by designers is to add a slightly greater quantity of molecular sieve 13X than calculated for normal amounts of contaminants.
5.4.3
LIQUID OXYGEN PURGE
The practice of maintaining a small, continuous LOX purge from the main condenser (as has been standard with Revex systems) is questionable. Purge rates of 0.5% are effective only when the contaminant concentration in the feed is low. If concentrations are high, the time needed to decrease their levels to acceptable values can be very long, because of the large volume of LOX in the main condenser compared with the volume of the purge. Using an auxiliary or pressure vaporizer and a LOX guard filter, however, enables LOX product to be drained from the main condenser only in a minimum continuous stream, or intermittently as needed. This will save substantial amounts of energy in plants of 1000 t=d or more of oxygen. In all cases, however, close visual monitoring by operators at every 4 h is strongly recommended. The increased reliability and safety of air separation plants using an FEP system should not lull operating departments into a false sense of security. Concentrations of CO2 and other hydrocarbons are continually rising in urban and industrial areas, so that designers are now using 400 vppm, and even 500 vppm, as the CO2 specification in heavily polluted regions. Some refineries and petrochemical plants now exhibit atmospheric concentrations of ethane and propane of 2 vppm and higher. Whenever this is the case a LOX guard filter is mandatory. In view of the dangerous, albeit low, concentrations of contaminants present, hydrocarbons should be monitored every 8 h, or at the very least, daily. If a total hydrocarbon analyzer is used, limits are sometimes set at 5–50 vppm in air and 40–150 vppm in LOX. It is far wiser and safer, however, to use an in-line gas chromatograph to check individual hydrocarbons, since a total hydrocarbon analyzer can promote a costly and unnecessary plant shutdown in the presence of high methane concentrations. Recently, a new methane–nonmethane hydrocarbon analyzer has arrived on the market. This apparatus employs a flame ionization detector (FID)
ß 2006 by Taylor & Francis Group, LLC.
TABLE 5.9 Recommended Procedures for Analyzers Trace CO2 in purified air Hydrocarbons in LOX
Infrared (1–10; 10–100; 100–1000 ppm) In-line gas chromatograph With the following points to be analyzed every 4 or 8 h: rich liquid at lower (HP) column; main condenser=reboiler; upstream of LOX guard filter; downstream of LOX guard filter
The following analyzers are recommended to follow the progression of pollutants though the entire system. Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
as the sensing element and provides a digital readout of either the methane component or the nonmethane components, which are the chief worry of oxygen plant operators. The cost of this unit is somewhat similar in cost to a standard total hydrocarbon analyzer.
5.4.4 ANALYZERS The in-line gas chromatograph should include an alarm for any abnormal high content of the following hydrocarbons: methane, carbon dioxide, ethane, ethylene, acetylene, propane, propylene, butane, butadiene, and acetone. Caveat, unless the instrument is specified and tested for accuracy, any readings under 50 vppb should be debatable (Table 5.9).
5.5 ACTIVATED ALUMINAS FOR FRONT-END PURIFICATION SYSTEMS 5.5.1 BACKGROUND As far back as 1957, the research department of L’Air Liquide in Paris reported in a private in-house communication that during tests for the desiccation of process air with an activated alumina supplied by Pechiney SA (France), it was observed that the absorbent also had removed a fair quantity of carbon dioxide. At the time, however, the synthetic zeolites had come into vogue for contaminant removal, and any further testing for carbon dioxide adsorption with activated alumina was abandoned. Then in 1984, Goodboy and Fleming13 of the Aluminum Corporation of America published a paper on the possibilities of using modified aluminas for selective separation purposes by physical adsorption, in addition to their conventional role as desiccants. It should be understood that activated alumina has been defined as commercially available alumina, which may normally contain small quantities of combined ferric, sodium, and other oxides. The increased understanding of porosity control in the manufacture of activated aluminas has permitted producers not only to engineer pore size distribution, but also the degree of porosity in the adsorbent. Presently, manufacturers can supply various formulated aluminas with properties of selective adsorption, which can meet most specific process requirements of the industry. Moreover, certain commercially available aluminas are formulated to contain specific quantities of other materials to enhance their capacity for selective adsorption as well as other beneficial properties. There are various ways of treating aluminas to increase their capacity for adsorption. It can be done by calcining the adsorbent with an alkali metal oxide, or by impregnation with a decomposable salt at calcining temperatures. There are numerous patents on these processes.
ß 2006 by Taylor & Francis Group, LLC.
The salts include hydroxides, bicarbonates, phosphates, and organic acid salts, with a preference for potassium carbonate. One may ask why there is such activity in formulating aluminas to enhance their capacity for adsorption, when zeolites, both natural and synthetic, are already available on the market and have had a long history of successful industrial application. The answer is very simple. Formulating aluminas to improve their affinity for carbon dioxide and other contaminants such as hydrocarbons offers a strong commercial challenge to synthetic zeolites in terms of energy savings in pressure swing adsorption (PSA) as it is known14–16.
5.5.2
PRESSURE SWING ADSORPTION
With the development of different types of formulated alumina adsorbents it is now possible to remove contaminants such as water, carbon dioxide, and dangerous hydrocarbons at close to ambient temperatures, say 300 K. This possibility is of high interest because it eliminates not only the requirements for a precooling system to reduce ambient air temperature down to 278 K, or 283 K, but also the need for extra energy to heat the regeneration gas (waste nitrogen) up to 363 K and even to 423 K, to regenerate the adsorbent bed when using a zeolite of the 13X variety. The process used for this type of adsorption is named PSA, in contrast to temperature swing adsorption (TSA) for the prepurification system using an adsorbent such as a zeolite 13X, which has a greater capacity for adsorption, but requires extra heat for its regeneration. Actually, when one uses a zeolite 13X adsorbent, operating pressure is also reduced to atmospheric before regeneration heat is applied. So that in reality, the latter system should really be called pressure–temperature swing adsorption (PTSA), according to D.M. Ruthven in his Principles of Adsorption and Adsorption Processes. Nevertheless, returning to the PSA system, the latter operates at normal process air pressure, of say, 700 kPa (7 bar). After process air passes through a standard compressor aftercooler, either a shell and tube or a DCA water wash tower, to return as closely as possible to an ambient temperature it is sent to the PSA unit. Depending on the specific design of the PSA system, two or three vessels are employed, either in vertical or horizontal configuration. The vessels are filled with a single bed or multiple beds of the appropriate selective-modified aluminas for the removal of the desired contaminants. Impurities entrained by the process air are retained by the absorbent at the same temperature and pressure as the inlet process air. In this regard, the operation of the PSA is similar to the TSA system, except that no extra chilling is required. A standard shell and tube aftercooler after the main air compressor, or a DCA will suffice. As soon as the first bed becomes saturated, the switch valves swing over, and process air is passed to a second vessel for treatment. The first vessel is then depressurized, and the impurities eluted by dry waste nitrogen. The complete regeneration involves the production of a very low partial pressure, always at ambient temperature, but at a lower pressure. The low partial pressure is obtained by dropping pressure in the vessel containing the adsorbent down to atmospheric, sometimes even below (using a vacuum pump), and by sweeping a waste gas (dry nitrogen) over the adsorbent. Precooling the air in a TSA system down to say 283 K is carried out to reduce the quantity of water entrainment from the DCA unit, to increase the efficiency of adsorption of the molecular sieve, and to reduce the quantity of water vapor. In the case of the PSA system operating at a higher temperature, more water is entrained increasing the quantity of activated alumina required in the vessel. As the latter adsorbent has a lower initial unit price, this saving may offset the total purchase cost of the adsorbent system.
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Activated alumina has only about a 20% capacity to remove carbon dioxide, as compared with synthetic 13X zeolite, but it is said that the former has the advantage in that its effective capacity, that is the difference between adsorption and regeneration step loading, is just as high as 13X zeolite in PSA operations. Moreover, activated alumina adsorbs one-third less process air than the 13X zeolite, and that means a lower loss of energy in the desorption phase. Add the fact that the regeneration of water is an endothermic reaction, which lowers the temperature of regeneration. It should be borne in mind, however, that the use of a zeolite for water adsorption is not viable. The present practice in the design for TSA systems is to use a bed of activated alumina upstream for this purpose, and a 13X zeolite for the second bed downstream for the adsorption of carbon dioxide, and other contaminants such as dangerous hydrocarbons. In view of the temperature limitations of the regeneration cycle for the PSA system, the quantity of formulated alumina adsorbents is quite high. Therefore, in order to save on adsorbents, the total NEMA time cycle including adsorption and regeneration is very short, in fact at times, much less than 30 min. According to the proponents of the PSA system, there may be an advantage in this because the thermal effects due to the release of heat during adsorption and the consumption of heat during desorption are minimized. Moreover, short time cycles may lead to quicker start-ups. But these possible advantages are debatable. The very short time cycle of the switch valves in comparison with the TSA system of 4–8 h also means a higher cost in maintenance or replacement of the valves, whose switching action may now be compared with the Revex system. Another point to consider is that when formulated aluminas may adsorb one-third less feed air than the synthetic 13X zeolite, the high number of switching cycles compared with the TSA may also mean a higher loss of total feed air from blowdowns over a longer period of time. Furthermore, the thermal effects previously mentioned have been minimized enormously with the addition of activated alumina upstream for the almost complete removal of water, which is the prime culprit for the thermal phenomena. In fact, the quantity of activated alumina used for this purpose reached almost 90% of the molecular sieve by weight in one specific TSA application.
5.5.3 INDUSTRIAL APPLICATIONS
IN
AIR SEPARATION PLANTS
During the past few years, certain types of formulated aluminas developed commercially for selective adsorption have been used as prepurifying agents in air separation plants, for either nitrogen or oxygen production. The advantage of this application is that, when the affinity for physical adsorption of these aluminas may not be as strong as the comparable synthetic zeolites, the former can be used at normal ambient temperatures for both adsorption and desorption. No extra energy is required for regeneration. A fair number of air separation plants have already been in successful operation using the PSA system as a prepurifier16,17. According to recent publications, some of these new formulated aluminas, such as H-156 and Selexsorb CDX manufactured by ALCOA, outperform the molecular sieve 13X in carbon dioxide adsorption in PSA–VSA (vacuum swing adsorption) applications. Whereas the synthetic 13X has a higher equilibrium CO2 capacity, the carbon dioxide is more easily desorbed from Selexsorb CDX17,18. No heat is required for regeneration. Just a swing from operating to ambient pressure, and a simple sweep by a purge gas at low pressure, are sufficient to elute the undesirable pollutants. This results in an improved capacity of the system, which can then be made smaller or operated on a longer cycle. The operating sequence of a proposed system may be as follows: 1. On line 10–25 min 2. Depressurization, all the gaseous contents are vented
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3. Purge phase, adsorbed pollutants are eluted by the sweep gas 4. Repressurization, a controlled stream of prepurified process gas is used to repressurize the desorbed vessel Under stated conditions for a two-vessel configuration overall valve switching loss is about 2%, which is rather high. It is considered prudent, therefore, to use a three-vertical vessel design followed by a pressure equalization procedure. In other words, the vessel to be purged is equalized in pressure with the vessel that has completed its purge phase. This latter step is then followed by a complete depressurization phase, and a complete purge. Finally, the sequence follows a pressure equalization, at one-half the pressure and a complete repressurization. In many air plants, the three-vertical design has more often been employed. It is also recommended, however, to install a hot-wired backup electric heater to be used in cases of malfunction or underperformance.
5.5.4
OBSERVATIONS
ON
PSA PREPURIFICATION
According to an article reviewing PSA prepurification for air separation plants published in Chemical Engineering of April 1999, and coauthored by Notaro, Ackley, and Smolack of Praxair, the authors noted the following comments19: ‘‘Prepurification unit adsorbents are often regenerated by reducing pressure alone, and without heating the purge gas stream, at air separation plants in which 30%–50% of the feed is available as purge gas. Examples include plants designed to deliver only oxygen or only nitrogen as products, or in total products plants where only part of the nitrogen is delivered as product.’’
In PSA regeneration the working CO2 capacity of the adsorbent is significantly reduced, particularly with zeolite adsorbents, from which desorption of CO2 is difficult due to their strong affinity. This reduced capacity would normally dictate a larger adsorber if regeneration took place at typical frequencies (2–7 h) for adsorption-step times in TSA prepurifiers. Adsorbent efficiency can instead be raised by the use of activated alumina as the predominant adsorbent, together with a six- or tenfold increase in the frequency of regeneration (20–40 min for adsorbent-step times). As a result, PSA prepurifiers may often be the more cost-effective solution for contaminant removal in small air separation plants. Removal of acetylene from feed streams, however, is a key issue in the design of PSA prepurifiers. In TSA prepurifiers the presence of a zeolite assures that CO2 breakthrough occurs ahead of acetylene breakthrough10. In PSA prepurifiers using only alumina the order is reversed20. Toward the end of the adsorption run, therefore, some acetylene may pass through the unadsorbed. These issues have been addressed by novel designs based on the use of adsorbent layering20,21 and mixtures22, especially with the use of aluminas formulated for this purpose. In addition to acetylene, one should also keep in mind the effects of ethane, ethylene, and especially propane, which are also dangerous contaminants and are not totally removed even by zeolites. The PSA system has always used the multiple bed (layer) design. The upstream layer uses activated alumina to remove water similar to the TSA system, but the PSA uses a much higher quantity of the material. In the downstream layer, the PSA system uses formulated aluminas for the removal of carbon dioxide and other dangerous hydrocarbons. This is also similar to the TSA system, but the latter uses a synthetic zeolite (13X), which has a higher affinity for the contaminants, and consequently requires a higher energy for its regeneration23,24.
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In comparison to TSA prepurification, the faster time cycle of PSA processes results in a greater loss of compressed process air feed during depressurization (blowdown losses). The higher required purge flow also generates a greater pressure drop across the adsorbent. These additional losses may be offset (particularly for smaller plants) by eliminating purge heating equipment and related energy requirements, and by reducing the number of stages for aftercooling feed process air.
5.5.5 FIELD OBSERVATIONS In an actual study of an oxygen plant with a process air requirement of 40,000 Nm3=h the PSA system shows some definite disadvantages. It requires a larger quantity of adsorbent, and a high flow of sweep gas for regeneration (Table 5.10). The PSA unit proposed used only two vertical vessels. If the design involved three vertical vessels, as most larger plants require, the cost for equipment, adsorbents, and switching valves would be much higher, not to mention the additional operating cost for valve maintenance and replacement. The PSA system for the prepurification of process air has been in use with relatively good success in a number of plants ranging from small capacities up to a 1500 t=d oxygen plant. These units have been designed to treat process air even at a temperature of 323 K. The application of a newly developed selective adsorption aluminas has thus challenged molecular sieves in the field of prepurification of process air in some areas, especially those in which total product recovery is low, and the cost of energy to regenerate adsorbents is very high. In the long term, however, one has to examine very carefully both the advantages and disadvantages of each material for the prepurification of process air25–27. For example, in the choice of aluminas, the use of short switching time cycles of say, 12 min or 43,000 per year if three vessels are used, may lead to very high mechanical maintenance as well as high replacement cost for switching valves. Assuming an equipment design for an air temperature of 323 K to assure a good performance, thus may also lead to an overdesign in terms of material, as well as increased pressure drop. Also with synthetic zeolites such as 13X, its high affinity for contaminants may lead to a higher regeneration energy use. To summarize, both selective adsorption aluminas and synthetic zeolites have a definite place in the field of prepurification. Individual cases will determine specific applications.
TABLE 5.10 Comparison in Use of PSA and TSA (PTSA) in an Actual Operating Test Items NEMA cycle (min) Adsorb temperature (K) Regeneration temperature (K) Heater Evaporative cooler Pressure drop (kPa) Adsorbent(s) (kg) Kilogram of adsorb=K Nm3=h of process air= hour of adsorption Valve switching per year
PSA
TSA (PTSA)
28 303 303 N=R N=R 17.22 12,600 4,010
130 283 393 Required Required 6.89 1,900 113 (including alumina)
36,428
1,700
Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
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REFERENCES 1. Kerry, F.G. Inc., 2006. 2. Breck, D.W. 1974. Zeolite molecular sieves, ed. E. Robert, 3–16. Malabar, FL: Krieger Publishing Company. 3. ALCOA Bulletin F-35-14480. 4. Breck, D.W. 1974. Crystalline molecular sieves. J Chem Edu 4P:678–690. 5. UOP Pamphlet F.1979J-2M-1=90. 6. Kerry, F.G. 1991. Front-ends for air separation plants—The cold facts. Chem Eng Prog (August):48–54. 7. Ruthven, D.N. 1984. Principles of adsorption and adsorption processes. New York: John Wiley & Sons. 8. Grenier, M., J.Y. Lehman, P. Petit, and D. Eyre. 1984. Adsorption purification for air separation units. Cryogenic processes and equipment. Book No GOO283. New York: American Society for Mechanical Engineers. 9. Isalski, W.H. 1989. Separation of gases, 216. Oxford: Oxford Science Publications. 10. Sefton, V.B. 1962. Contaminants: Source, physical and chemical properties, basis of limits, presented at Compressed Gas Association Symposium on air Separation Plant Safety, October 5, New York. 11. Reyhing, J. 1983. Removing hydrocarbons from the process air of separation plants using molecular sieve adsorbers. Linde Reports on Science and Technology 36:14. 12. Ermenc, E.D. 1956. Wisconsin process system for recovery of dilute oxides of nitrogen. Chem Eng Prog 52:488. 13. Karwat, E. 1961. Hydrocarbon control in air separators. Chem Eng Prog 57:41. 14. Goodboy, K.P., and H.L. Fleming. 1984. Trends in adsorption with aluminas. Chem Eng Prog 80:63. 15. Keller, G.E. 1995. Adsorption building on a solid foundation. Chem Eng Prog 91:56–67. 16. Smith, D.L. 1996. Optimizing solid bed adsorption systems. Hydrocarb Process:129–132. 17. Kalbassi, M.A. 1996. Advanced process swing adsorption (PSA) air purification systems. International Institute of Refrigeration, Meeting on air Separation Technology (MUST’96), 10–12 October 1996, Munich, Germany, pp. 159–173. 18. Golden, T.C., and M. Kalbassi. 1998. Thermally enhanced pressure swing adsorption (TEPSA) air pre-purification system. AIChE Annual Meeting. Miami, FL, 15–20 November 1998. 19. Smith, D.L. 1999. Optimization of air plant pre-purifiers through use of combination adsorbent beds. Seminario Ass. Brasileira de Metais e Materiais, XIV Encontro de Produtores e Consumidores de Gases Industriais. 14–16 Setembro 17–23. Osasco, SP. Brasil. 20. Notaro, F., M.W. Ackley, and J. Smolack. 1999. Recover industrial gases via adsorption. Chem Eng:104–108. 21. Notaro, F. 1996. Advances in ambient temperature air separation. International Institute of Refrigeration, Meeting on Air Separation Technology, Munich, Germany, 10–11 October. 22. Jain, R. Carbon dioxide removal by pressure swing adsorption in an alumina adsorption bed. U.S. Patent 5,232,474, August 3, 1993. 23. Leavitt, F.W. A PSA pre-purifier for the removal of contaminants in a feed gas stream. U.S. Patent 5,769,928, June 23, 1998. 24. Golden, T.C., et al. A process for the adsorption of at least carbon dioxide, water, and oxides of nitrogen and preferably acetylene from a feed gas. U.S. Patent 5,779,767, July 14, 1998. 25. Industrial Gases 1997 Report MA28C (’97)-1. U.S. Department of Commerce, Bureau of Census, September 1998. 26. LaCava, A.I., A.I. Shirley, and R. Ramachandran. 1998. How to specify pressure swing adsorption units? Chem Eng 105(6):110–118. 27. Knaebel, K.S. 1999. The basics of adsorber design. Chem Eng Prog 106(4):92–101.
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FURTHER READING ON THE SUBJECT OF ADSORPTION AND CARBON DIOXIDE BUILD-UP Norobilski, J.J., and J.S. Schneider. Particle loader. U.S. Patent 5,324,159, June 8, 1994. Kumar, R. 1996. Vacuum swing adsorption for oxygen production—A historical perspective. Separ Sci Technol 31:877–893. Ackley, M.W., et al. Multiple adsorbent loading method and apparatus for a radial flow adsorber. U.S. Patent 5,836,362, November 17, 1998. Houghton, J.T., et al. 1996. Climate change 1995. The science of climate change. Cambridge: Cambridge University Press. Case grows for climate change. Chem Eng News, August 9, 1999, p. 16.
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6
Product Liquefaction, Storage, and Transportation
6.1 BACKGROUND Early air separation plants were designed primarily for the economic production of low-cost gaseous products to be supplied in high-pressure cylinders (150 bar) for industrial applications. The production of either liquid oxygen or liquid nitrogen was of a secondary consideration. Liquids were produced, however, but in small quantities, and essentially for use in research work at low temperatures in university laboratories. Small plants for total liquid production were also designed for the supply of liquid oxygen to coal mines in eastern France and in certain locations in the United States for the manufacture of liquid oxygen explosives. This application involved the use of activated carbon saturated with liquid oxygen enclosed in cartridges of nonflammable material. It has now been completely eliminated because of the relative high cost of cryogenic liquids versus the lower cost of standard chemical explosives. During World War II, however, when the availability of high-pressure industrial gas cylinders became very scarce because of the higher priority for forgings for the military, the concept of large-scale production and distribution of products in a liquid phase began to take form. Whereas the production cost of liquids was high in terms of energy consumption, this was more than offset by the lower transportation costs for the equivalent quantity of product. With the use of well-designed double wall vacuum-insulated storage tanks, liquefied gases were transported across the continent by trucks and railroad tank cars with commercially acceptable evaporation losses. The large-scale production and distribution of liquid products also initiated the development of more efficient liquefaction cycles, the design of larger liquid product storage tanks, a more extensive distribution of liquid products, and the newer applications of liquid cryogens. Oxygen, nitrogen, argon, hydrogen, and even helium have been liquefied and transported by railroad tank cars and large trucks, the latter hauling up to 20 t (liquid oxygen) at a time. Even relatively small individual customers are now supplied by the use of small double wall vacuum-insulated containers holding up to 265 L and more of cryogens. To provide these services, cryogenic producers are using large cryogenic storage facilities with individual capacities of up to 2500 t (2200 m3), and in one case even up to 5000 t.
6.1.1 NEW APPLICATIONS 1. Stored liquids are used to supply products to a process for peak load-saving applications, during an emergency shutdown of the main air separation plant, or during periods of planned maintenance. For example, in partial oxidation for the production of synthesis gas, if the oxyfuel burner is extinguished for some reason for an interval of 10 s or more, it may lead to a total plant shutdown for at least 2 or 3 d. This is economically unacceptable. For this reason, a small high-pressure storage tank of liquid
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Bar graph showing relationship between power consumption per tonne of LOX at various production rates (the production is only from the cold box)
Power consumption, KWh/t of LOX
1000
800
600
400
200
20
40
60
80
100
Cold box production of LOX, t/d
FIGURE 6.1 Bar graph showing relationship between power consumption per tonne of LOX at various production rates (the production is only from the cold box). (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
oxygen is always connected directly to the main process oxygen pipeline. It is ready to supply oxygen instantaneously at the required operating pressure for at least 30 min while awaiting the start-up and operation of the stand-by liquid oxygen pumps, and vaporizers drawing product LOX from the main storage tank(s) (see Section 6.3.5.3, Figure 6.1). 2. To maintain the quality of food after preparation, cooling must be fast. Liquid nitrogen can be used to cool down perishable food at a temperature of 223 K very quickly, and to maintain this temperature at a steady level. 3. Liquid helium can be used for the refrigeration of processes that require a temperature of 13 K, such as MRI equipment in hospitals. 4. Liquid nitrogen can be used for cryosurgery. At liquid nitrogen temperatures of 77.3 K, the surrounding tissue is cauterized, thus preventing any abnormal cells from entering the bloodstream. Liquid nitrogen is also used for freezing human and animal sperm, which can be held alive for up to 20 years.
6.2
PRODUCT LIQUEFACTION
The production of pure liquid products from an air separation plant is possible either by direct extraction in small quantities or by the use of auxiliary refrigeration equipment for the liquefaction of large quantities. The process is fairly simple, as long as the designer obeys the second law of thermodynamics. As the extraction of pure products in a liquid phase increases, there is an equivalent loss of thermal energy, or enthalpy. In order to maintain a correct energy (enthalpy) balance in the process, therefore, an equivalent increase of external energy must enter the process. There is another factor (i.e., enthalpy balance), moreover, that
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has been mentioned in a previous chapter 3, but not always appreciated or understood by engineers inexperienced in cryogenic technology.
6.2.1 ENTHALPY BALANCE In air separation plants that have been designed for low-cost gaseous products, process air is refrigerated, liquefied, fractionated, and the individual gaseous products are revaporized. The refrigeration capacity of the liquefied air is almost completely recovered in a countercurrent heat exchanger system within the cryogenic process equipment. Within 5 K, the gaseous products leaving the primary heat exchangers have approximately the same total enthalpy as the incoming air. The overall discrepancy in enthalpy is due to the heat gain from outside of the cold box, and by imperfect heat exchange. The enthalpy is returned into balance, however, by the generation of extra refrigeration derived from the expansion machine. The extraction of any liquid product, therefore, whether it be oxygen, nitrogen, or argon, in any quantity and combination requires additional external refrigeration equipment depending on the total quantity of liquid products extracted.
6.2.2 DIRECT EXTRACTION As noted, using a standard air separation plant for gaseous production, it is possible to extract a small quantity of liquid products because the addition and operation of the turboexpander may have the design capacity to generate enough additional refrigeration to offset the enthalpy loss from a small liquid extraction. This procedure involves extracting a small quantity of process air mixture from the high-pressure column, sending it, first through the primary heat exchangers in order to preheat it, and then to the expansion machine to generate additional refrigeration capacity. This procedure, however, is limited by the quantity of surplus heat exchange surface, if any, that may be available in the primary heat exchanger system. This surplus heat is necessary to preheat the extra air–nitrogen mixture extracted from the high-pressure column, and still maintain a correct heat balance within the primary heat exchanger system, thus avoiding opening up of the warm end temperature difference in the primary heat exchangers, and thus a loss of refrigeration. In a conventional process cycle, it is not possible to liquefy and extract more than 2% of the total process air feed without the addition of extra equipment. Increased turbine flow has another disadvantage in that it adversely affects the rectification of air in the upper (lowpressure) column, reducing the desired recovery of oxygen and argon. As explained earlier, there is no sufficient heat exchange surplus area in the primary heat exchangers in a standard air separation plant, to preheat the air–nitrogen fraction entering the turboexpander in order to compensate for the additional refrigeration required. Nevertheless, the additional liquefaction of 2% air feed represents an additional 10% of liquid oxygen production, which is more than enough in many cases. In the latter case, moreover, one may have to consider the use of two expansion turbines operating in parallel. Springman1 provided, in his classic article on liquefaction, the relationship between energy and air liquefaction for direct liquid oxygen extraction (Table 6.1). In actual practice and for a low-pressure plant with a capacity of 1000 t=d of oxygen, the total energy consumption per tonne of product has been found to be significantly higher than the one given in Table 6.1. For a liquid oxygen production of about 1% of the total, the net energy was 400 kWh=t, for a production of about 3% it was around 800 kWh=t, and for 5%–10% it was about 900 kWh=t. The net energy also includes all losses due to evaporation and pumping from the cold box to the storage tank (Table 6.2). According to Springman’s calculation, the energy cost for liquid oxygen in the aforementioned process mode is commercially
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TABLE 6.1 Conclusion of Springman in Direct LOX Extraction from Air Separation Up to 0.6% of liquefied air Up to 1.0% of liquefied air Up to 2.02% of liquefied air
0.70 kWh=Nm3=h 0.85 kWh=Nm3=h 1.00 kWh=Nm3=h
490 kWh=t O2 595 kWh=t O2 700 kWh=t O2
Source: From Springmann, H., Chem. Eng., May 13, 58, 1985. With permission.
unacceptable when one reaches a production level of 5% of the total oxygen production capacity of the plant. During this liquid production, moreover, there is a simultaneous loss of gaseous oxygen production of approximately three times the liquid equivalent in tonnes. Although this high-energy cost may be viable in some rare cases, it is completely unacceptable
TABLE 6.2 Performance Tests 1200 ST=D Oxygen Plant Mode D 1 GOX 1 Maintenance LOX Level Duration of test Stability of operation Process adjustments Average airflow Average GOX production or Average LOX production (net) Total losses (LOX) 2500 t storage tank Small storage tanks Cooling and bumping cryopumps Total losses Total LOX produced Total O2 produced (c þ f) Net O2 produced (c þ d) Power consumed Energy=net tonne (i 24 h)=h Energy=gross tonne (i 24 h)=g Process efficiency (airflow) 6456.117 0.2095 O2 produced 1210.88 0.99526 Gas equivalent LOX ¼ 20.1 1.35 Total oxygen recovery Total oxygen loss (j k) Process efficiency (k=j) Waste nitrogen (a k) Calculated (l=n) Analyzed at plant
26 h Excellent Minor optimization 6456.117 KSCFH 1210.88 KSCFH 1202.96 ST=D 1.0 ST=D
(a) (b) (c) (d)
10.0 t=d 3.1 t=d 6.0 t=d 19.1 t=d 20.1 t=d 1223.06 ST=D 1203.96 ST=D 18,771.66 kWh=h 374.2 kW=t 368.355 kW=t
(e) (f) (g) (h) (i) (A) (B)
1352.556 KSCFH oxygen 1205.140 KSCFH oxygen 27.135 KSCFH oxygen 1232.275 KSCFH oxygen 120.281 KSCFH 91.11% 5223.842 KSCFH 2.3025% 2.31%
(j)
(k) (l) (m) (n) (o) (p)
N.B. Comparing (o) with (p) it is quite obvious that the plant is producing a fair tonnage of LOX (21.1) of which (20.1) unfortunately is used up to maintain the liquid level at all the storage tanks, also to keep the cryopumps chilled as well as in a ready and operable state. Source: Courtesy of F.G. Kerry, Inc., 2006. With permission.
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Low-pressure nitrogen Warm exchanger
Feed compressor
Medium-pressure nitrogen
Cold exchanger
LIN to storage Separator Subcooler
Liquefier exchanger Recycle compressor
Turbo blower
Turbo expander
LAIN assist to ASU Pressurized nitrogen from ASU
FIGURE 6.2 Liquefaction of nitrogen from independent liquefier. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
for larger requirements of liquid products. It is obvious, therefore, that other more efficient liquefaction process systems have to be employed (see Table 6.2 and Figure 6.2).
6.2.3 BASIC DESIGN PARAMETERS
FOR AN
EFFICIENT LIQUEFACTION SYSTEM
Between the direct extraction of liquid product from a standard low-pressure process cycle and the classic high-pressure total liquid products plant, there is a multitude of process cycle selections depending on the imagination of the process designer and the market requirements. This chapter will limit itself to a few process cycles that are in wide use. The standard total liquid products and high-pressure cycle reached its peak during the 1950s and 1960s with a top overall capacity of around 100 t=d, using a reciprocating expansion engine with glass-impregnated Teflon rings (see Chapter 10.3). This system also employed an independent nitrogen recycle compressor in order to produce simultaneously high-purity product nitrogen. With the development of brazed aluminum heat exchangers that could withstand high pressures and the availability of highly efficient radial expansion turbines, however, the high-pressure system was quickly replaced with low-pressure auxiliary liquid-producing systems (so-called liquefiers), with very low-energy requirements and reasonable investment costs. An energy-efficient liquefier for the large-scale production of liquid products involves a separate system with its own closed refrigeration cycle. It includes a compression and an expansion system of its own. The refrigeration fluid used for compression and expansion is generally gaseous nitrogen, used because of its inertness, and therefore safety, even at high pressures (100 bar). The liquefier unit receives pure gaseous products from the nearby air separation plant, and converts them into a liquid phase. With the use of multipass, high efficiency heat exchangers, the system can liquefy independently any combination of products, within the preset refrigeration capacity of the liquefier. Standard process equipment for the liquefaction of gaseous products into the liquid phase is simple enough. The process involves a warm exchanger system wherein the temperature is lowered close to 176 K, a cold exchanger system that lowers the temperature still further close to 77.3 K, and a liquefier vessel. It is also prudent to add a separator and a liquid
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product subcooler, in order to drop the temperature of the liquid produced by a few degrees, so as to minimize the evaporation and flash losses as the liquid is transferred to an outside storage tank.
6.2.4 TYPES OF LIQUEFIERS There are two general classes of liquefiers: separate or independent type, and the integrated type, which is tied to the process of an air separation plant. 6.2.4.1
Independent Liquefier (Figure 6.2)
Where the product liquefier is to be applied to an existing air separation plant, or where the owner and the operator wish the two process systems to be completely independent of each other, the liquefaction cycle involves an external refrigeration system, which will be completely independent of that of the air separation unit. The external refrigeration cycle proposed for a standard independent liquefaction system is as a rule of the closed recycle loop type. The medium used is nitrogen, which can be handled and compressed to high pressures safely. Nitrogen has a boiling point of 77.3 K at 101.325 kPa, and can be used as a refrigerant for liquefying almost all gases including hydrogen. Two compressors operating at two separate compression levels are used for the main refrigeration cycle. The first compression level is around 6.2 bar, which is very convenient, because it is somewhat close to the operating pressure of the lower (high-pressure) column. Gaseous nitrogen can be withdrawn directly from the high-pressure column and used in the second compression level of the liquefier. The second compression level is around 31 bar, which is also convenient, because it involves a standard radial compressor commonly found in the market. Further compression of the nitrogen refrigerant stream can also be achieved with the use of a booster compressor (turboblower) mounted on the same shaft as the expansion turbine. After a predetermined pressure level is reached, the high-pressure nitrogen is expanded and returned through the exchanger system providing refrigeration to the incoming gaseous product streams to be liquefied. After the expanded nitrogen refrigerant reaches the lower level of around 6.2 bar, and releases its refrigeration into the heat exchangers, it is recycled to the inlet of the highpressure level compressor. It may be useful to provide a small separate external mechanical refrigeration unit (MRU), using an environmentally acceptable refrigerant, and located between the warm and the cold exchangers, because it will assist in the overall conservation of energy. Such units operate between 233 and 216 K (408C) and can save up to 20% in the overall energy of the system. As can be seen from Figure 6.2, the process is fairly simple. The low-pressure nitrogen refrigerant from the low-pressure column LP and the warm exchanger WHX is fed into the medium-pressure recycle compressor FC, and boosted to 6.2 bar. This total stream is joined by a medium-pressure recycle stream also from the warm exchanger WHX, and both are fed into the high-pressure recycle compressor REC where they are boosted to a higher pressure of around 31 bar, as previously noted. This pressure level may also be raised to a higher level with the use of a turboblower mounted on the same shaft as the expansion turbine. Once the nitrogen refrigerant reaches its final pressure, it is passed first through the WHX, then through the cold exchanger CHX, then to the liquefier LX, and finally into a separator SEP and a subcooler SBC. In between the warm and cold exchangers, the designer may use an MRU. If the oxygen product has to be liquefied simultaneously, the same equipment is involved, but an additional pass is added to the exchanger and a separate oxygen liquefier–separator is included, as can be seen in Figure 6.3. In this figure, high-pressure gaseous oxygen HPO traverses the heat exchanger system of the liquefaction system, undergoing a progressive drop
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High-pressure oxygen
Low-pressure nitrogen Refrigeration machine
Feed compressor
Warm exchanger
Cold exchanger
Medium-pressure nitrogen
LIN to storage
Separator Liquefier exchanger
Recycle compressor
Turbo blower
Pressurized nitrogen from ASU
Subcooler
Turbo expander LIN assist to ASU LOX separator and subcooler LOX to storage tanks
FIGURE 6.3 Simultaneous liquefaction of O2 and N2 from independent liquefier. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
in temperature without any serious change in pressure. Finally, this oxygen stream enters its separate liquefier by its own coil. It undergoes condensation as well as subcooling by heat exchange from the vaporization of liquid nitrogen, the latter having been fed from the liquid nitrogen separator. The liquid oxygen is then transferred to an outside storage tank, whereas the vaporized liquid nitrogen is returned through the heat exchanger system to recover its refrigeration, and recycled to the high-pressure compressor, REC. The process is designed for a preset refrigeration capacity of the liquefier unit for the production of liquid nitrogen. If oxygen is liquefied exclusively, the final quantity of the product will be slightly less than the calculated liquid nitrogen product, because the liquefaction of oxygen requires a higher refrigeration load. With multipass exchangers it is possible to liquefy oxygen and nitrogen simultaneously, but the total production of both products should remain within the refrigeration limits of the designed nitrogen cycle. Adding an environmentally acceptable refrigerant, it will be possible to upgrade the refrigeration capacity of the system by about 20% to 25%. 6.2.4.2 Integrated Liquefier (Figure 6.4) Integrated liquefiers are very advantageous for merchant plants because they offer a lower overall investment cost to users, as well as a wide variety of operational modes involving both liquid and gaseous products. These liquefiers also offer an operational simplification as the plant has a single central control system. The liquefaction section of the total plant is similar to a separate liquefier. It consists of a warm exchanger, a cold exchanger, a liquefier, and a subcooler. In the integrated cycle, however, nitrogen refrigerant is supplied directly from the fractionation column of the air separation plant within the same cold box. This nitrogen stream is first used in the warm exchanger of the liquefier to precool the high-pressure nitrogen cycle, then discharged from the cold box and fed to the recycle compressor. In turn, the liquid nitrogen produced is
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Liquid oxygen Liquid nitrogen
Primary heat exchanger
Primary heat exchanger
Oxygen Waste nitrogen
Expansion turbine
Lowpressure column
Sub cooler
Main condenser
Air Nitrogen recycle Compressed nitrogen Highpressure column
Liquefier exchanger
Rich liquid External refrigeration
Expansion turbine
FIGURE 6.4 Gaseous oxygen plant with integrated liquefier. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
extracted and sent to the lower (high-pressure) column. There, a portion is withdrawn and passed through a subcooler before being sent to storage as liquid nitrogen product. A portion of the liquid nitrogen sent to the lower (high-pressure) column is used as reflux to allow withdrawal of oxygen in liquid form from the bottom of the main condenser and vaporizer. The liquefier section, therefore, must always be in operation in order to supply reflux to the low-pressure fractionation column even though no liquid nitrogen production may be required for outside storage. This latter problem may be eliminated, however, with the use of a separate turboexpander for the air fractionation unit. This solution is costly, however, and is really unnecessary in most cases. The other option during a low demand for liquid production is to supply refrigeration to the high-pressure column from the liquid nitrogen storage tank, which is known in the industry as a liquid-assist operation. All cryogenic equipment for both air fractionation and liquefaction is housed in a single cold box for a lower investment and convenience of maintenance. An earlier version of an integrated liquefier operating at comparatively low-pressure for the simultaneous liquefaction of oxygen and nitrogen is shown in Figure 6.4. In this particular flow sheet, nitrogen refrigeration cycle supplies liquid nitrogen to the top shelf of the lower (high-pressure) column, in exchange for pressurized gaseous nitrogen at 6.2 bar, also from the high-pressure column. With this procedure, one reduces the quantity of refrigeration available in the liquid oxygen at the main condenser, and is necessary for the condensation of gaseous nitrogen at the top shelf. Therefore, it is possible to withdraw liquid oxygen directly from the main condenser. In this process cycle, the recycled nitrogen is fed into the compressor at 6.2 bar, and the discharged nitrogen at 31 bar. As mentioned earlier for the independent liquefier, energy saving of about 20%–25% can be made with the use of an MRU operating with a difference of 40 K between the warm and the cold heat exchangers. With the
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integrated process cycle, as described in Figure 6.4, the designer can avoid compressing the pure oxygen feed, and eliminate the use of special passages in the heat exchangers. Ancillary equipment, i.e., manifold piping and general instrumentation, can also be simplified. One may also consider eliminating the expansion turbine from the air separation unit, as the liquefier unit can be designed to supply the entire refrigeration requirements for both the air separation and the liquefier units. It should be kept in mind, however, that this scheme may lead to a slight loss in enthalpy from the extraction of pressurized nitrogen from the high-pressure column, extra losses from the heat gain due to insulation inefficiencies at the cold box, and temperature differences of the outgoing products at the outlet of the primary heat exchangers. Nevertheless, the designer can compensate for these losses by making adjustments in the overall capacity of the liquefier, and in the process airflow for the air separation unit. 6.2.4.3 Very High-Pressure Liquefiers Since 1995, designers have achieved higher nitrogen recycle pressures, 60 and 70 bar, for the liquefaction of oxygen and nitrogen, for greater economies of energy. In Figure 6.5 it can be seen that low-pressure nitrogen makeup feed is supplied from the air plant and sent to the feed gas compressor FC, where it is raised to 6.5 bar. This stream is joined by another
After cooler
After cooler
Recycle compressor
Feed from LPC
Recycle compressor Warm heat exchanger Mechanical refrigeration
After cooler
Warm expansion Warm turbo booster
GAN from PHX GAN from PHX Cold exchanger
After cooler
Cold expansion Cold turbo booster
Separator
LIN to top shelf HPC GAN from PHX
77.3 K LIN to top shelf HPC
FIGURE 6.5 High-pressure integrated liquefier. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
medium-pressure stream also at 6.5 bar from the prime heat exchangers of the air plant. Both streams are now sent to the nitrogen recycle unit REC and raised to 31 bar, which after compression is led to the first high-pressure cold booster CB, where it is compressed to about 41.3 bar. Then this stream, after passing through its own aftercooler, is sent to a second blower WB, coupled to the warm expander WXT, and where it is further compressed to about 60 bar. After this final compression, the main nitrogen refrigerant stream passes through another standard aftercooler, and at its highest pressure, is led to the warm exchanger WHX, where it is precooled to about 161 K. This precooling is brought about by the countercurrent effects of a medium-pressure nitrogen stream supplied by the primary heat exchangers of the air plant, and by the expansion of part of the recycle stream from 31 bar down to 6.5 bar at the ‘‘warm’’ expansion turbine WXT. Both of the precooling streams are returned to the outside nitrogen recycle compressor REC, with the warm heat exchanger WHX. The high-pressure stream still at 60 bar, and at 161 K, continues to the cold heat exchanger CHX. Before its entry to the exchanger (liquefier), a part of it is sent to the expansion turbine CXT, where it is expanded and reaches its final liquefaction temperature of around 77.3 K. This liquid enters the separator SEP. In the separator the portion remaining as liquid is now subcooled and sent outside to the storage tanks. The portion in the separator, which is in vapor phase, is returned through the exchanger system as a refrigerant, and finally fed to the recycle compressor. Another portion of liquid nitrogen from the liquefier is sent to the top shelf of the lower (high-pressure) column to be used as reflux in the fractionation columns. The final portion of the liquid nitrogen product may be transferred to the pure argon column, if the latter is used, to purify the argon. Calculated efficiency in each turboexpander is 85%–86%, whereas the turbobooster efficiencies are approximately 75%. As in the case of the independent liquefier, both oxygen and nitrogen can be liquefied either separately or simultaneously, as long as the total refrigeration requirement does not surpass the maximum refrigeration capability of the nitrogen recycle. It should also be understood that the nitrogen recycle must be continuously supplied with fresh gaseous nitrogen feed: (a) to make up for the liquid nitrogen extracted from the plant and supplied to storage; (b) to supply liquid nitrogen as reflux to the fractionation column; and (c) to make up for any losses within the process system. 6.2.4.4
General Summary
With the exception of a few cases involving metallurgical and petrochemical operations in isolated areas, the prime use of liquefiers is for merchant plants that have to modify their mode of operation from day to day depending on the ever-changing local market for liquids and gases. If local market conditions require a fairly large production of liquids, say above 300 t=d, it is important to design a liquefier with the highest possible energy efficiency, because energy is the most important factor in production costs. For large liquefiers, therefore, it pays to add any additional equipment that may save operating energy. On the other hand, having liquid oxygen, nitrogen, or liquid just sitting idle in large storage tanks means a loss of revenue. Above the liquefaction of 300 t=d of combined products, it is wise to consider the use of high-pressure ratios for compressing nitrogen, if the latter is the major and the final refrigerant. It is also recommended to employ a multistage MRU in the process, as a form of cascading principle. This will reduce the energy consumption considerably. If MRU is used, the refrigerant should be environmentally acceptable, such as HFC 134a (1,1,1,2 tetrafluoroethane) or equivalent. The use of an ammonia refrigeration unit, while acceptable from a process standpoint, can be a headache in mechanical reliability, continual maintenance, and toxic odor.
ß 2006 by Taylor & Francis Group, LLC.
As for energy savings from high nitrogen compression, the use of pressures up to 70 bar and even higher is also recommended, but with the use of turboblowers directly coupled to the expansion turbines. For these very high pressures, one has to consider the choice between the availability of brazed aluminum and heli-wound coiled heat exchangers. The final choice for high pressures also depends on the availability and reliability of the turboblower and expansion turbine combination in one complete unit. Another possible thought is to use a liquid expansion turbine, which has shown some advantages in expansion from very high pressures, even up to 100 bar. 6.2.4.5 Energy Economics Some 30 years ago, energy consumption in production of liquid oxygen was approximately 600 kWh=t. Presently, the energy consumption has been reduced to 400 kWh=t or even less. If one considers the theoretical Carnot cycle, the power necessary to liquefy 1 t of nitrogen is around 80 kWh. Though unachievable, this gives the process designer a good target for the future.
6.3 CRYOGENIC STORAGE FACILITIES 6.3.1 GENERAL CONSIDERATIONS Liquid storage of products from an air plant is a viable procedure in order to coordinate the economic steady-state operation of the plant and the fluctuating demands of the market. These demands may range in capacity anywhere from a 1 L Dewar flask to a delivery requiring the use of a railroad tank car. Choosing a specific design for the cryogenic liquid storage tank or container, especially one of a very large size, requires a complete economic study. Inventory of cryogenic liquids represents a high fixed capital, as well as a high-energy operating cost. Therefore, a careful economic balance should be carried out on the desired inventory level, keeping in mind lost production due to a plant shutdown or loss of profit in sales to third parties. The design capacity of a cryogenic liquid storage system should be specified only slightly above the calculated inventory. In turn, inventory should be calculated to take care of: . . .
Peak demands of a process using the product Planned or emergency shutdowns Sale of product to third parties
6.3.2 GEOGRAPHIC CONSIDERATIONS 6.3.2.1 Ambient Temperature It is important to have a record of both the minimum and maximum temperatures, which occur in the area. This will affect the choice of materials, plate thickness, quantity, and quality of insulation. Daily variations in ambient temperature will affect the amount of thermal breathing, thus influencing the design of the vent recovery system, design pressure of the vessel, etc. The most economic storage temperature is one at which the vapor pressure of the contained liquid slightly exceeds the atmospheric pressure. 6.3.2.2 Wind Loading Considering the large surface of liquid storage tanks, it is necessary to specify the highest wind loads expected in the area. Under-specifying wind loads for reasons of economy could be potentially dangerous.
ß 2006 by Taylor & Francis Group, LLC.
6.3.2.3
Seismic Loadings2
In areas of high seismic conditions, tanks must be designed accordingly. The Uniform Building Code (UBC) Division IV Earthquake Design is recommended. The calculations are complex and should be entrusted only to an experienced engineer, or a group specializing in this type of work. 6.3.2.4
Soil Conditions and Land Cost
It is necessary to determine the allowable footing loads, and also the need of piling. Cost and availability of land to be used for the storage tanks will determine the optimum geometry of the storage tank. 6.3.2.5
Snow Loads
If snow is a possibility, the load should be considered in the design of the roof structure if one is necessary. Subfreezing weather will also affect the choice and design of ancillary equipments such as pumps, instrumentation, steam tracing, and general maintenance. 6.3.2.6
External Corrosion
Because of the high cost of the product, the potential hazard from spillage or overpressurization of its vapor, it is of the utmost importance to maintain the integrity of the storage insulation system by protecting the outer shell from corrosion. This requires the proper choice of materials, the use of suitable preparation, and painting procedures. The inner tank is generally made of either stainless steel, aluminum, or a 9% nickel steel alloy, so that no special procedures need to be followed, aside from maintaining a high standard for descaling, degreasing, and welding. 6.3.2.7
Availability and Dependability of Utilities
This will affect the choice and design of pumps, lighting, refrigeration and vaporization systems, instruments, steam tracing, rainwater runoff, and communications. 6.3.2.8
Local Neighborhood Characteristics
The characteristics of the local area where the tanks are to be erected must be reviewed, as this will influence design features for safety and the environment. In populated and industrial areas, elaborate safety features are necessary: fire fighting equipments, remote control valves, concrete drainage systems, dykes, etc. Needless to state, an initial public relation program in the selected area is important.
6.3.3 6.3.3.1
DESIGN PARAMETERS Low-Pressure Shop-Built Tanks
These are restricted to a capacity of approximately 80 m3, and follow API Specifications 620 recommended for Design and Construction of Large Welded Low-Pressure Storage Tanks. These rules allow the pressure in the vapor space above the liquid to reach, but not exceed, 987 mbar G. Materials of construction depend on the temperature level.
ß 2006 by Taylor & Francis Group, LLC.
6.3.3.2 Storage Vessels with Internal Pressure ASME Section VIII Division 1 or 2 gives design criteria for storage at pressures greater than 0.98 bar G. This code covers design, materials, fabrication, inspection, and testing of pressure vessels. The ASME Code is also an ANSI Standard, ANSI=ASME BPV—VIII-I. 6.3.3.3 Low-Pressure Field-Built Aluminum Tanks This design is covered by ANSI B96.1, Specifications for Aluminum Alloy Field-Erected Storage Tanks. The pressure limitation is close to atmospheric. 6.3.3.4 Flat Bottom Tanks These tanks have end enclosures designed for pressures near atmospheric. The shell must resist lifting forces arising from internal pressures. Hold down devices must permit the vessel to move readily in response to thermal displacements. This type of design provides cryogenic storage at a minimum capital cost, but special shapes may be dictated by service requirements: . . .
Pressure required for transfer of contents Evacuation (vacuum) of insulation space to reduce heat transfer Critical foundation conditions making other shapes or forms, such as spheres, more economical
The economic advantage of this type of design is that the inner container rests upon some form of load-bearing insulation, which carries the weight of the contents to the foundations. The bottom serves only as a seal, and is not subject to any significant stress. This permits it to be made of relatively thin material. 6.3.3.5 Spherical Containers A spherical inner tank enclosed in a spherical outer tank is geometrically the most economical for capacities up to 500 m3 as it offers a maximum internal volume for the minimum external surface. Nevertheless, it also demands a high cost of plate preparation and field erection. While it was a standard for the cryogenic industry over a long period of years, it has given way to the flat bottom shape, especially for very large storage volumes, for example, 1000 t and over for liquid oxygen. 6.3.3.6 Cylindrical Vessels (Horizontal or Vertical) This design is adaptable for internal pressures built up in cryogenic storage, originally designed for evacuated expanded perlite insulation, now redesigned for evacuated multilayer insulation. It is limited to small capacities because it is shop built as well as tested, and shipped as a complete unit, with all the ancillary equipment. They range in capacities from small liquid cylinders of 300 L to large 50,000 L (50 net tonnes) for customer on-site bulk stations. The cylindrical horizontal vessels with a capacity of approximately 20 t can also be mounted on trucks and transported to various clients.
6.3.4 DESIGN SELECTION For high-value products such as argon, hydrogen, and helium where the storage capacity is generally small, the obvious choice is the cylindrical vessel under pressure and with multilayer
ß 2006 by Taylor & Francis Group, LLC.
insulation with a high vacuum, in order to reduce boil-off losses. These can be reduced to 0.1% per day or even less. For more common but still expensive liquid products such as liquid oxygen and nitrogen, where the storage pressure of the vapor is slightly above atmospheric (up to 9.867 kPa G), the economic choice, the economic choice is limited to pressurized bulk vessels. (Figures 6.6 and Figure 6.7). When soil conditions are very poor, such that a flat bottom tank may require a large number of foundation piles, a spherical design should be evaluated as it may prove less costly. A flat bottom tank has a lower boil-off loss than a spherical shape in similar conditions of vapor pressure and insulation, with storage capacities in excess of 1000 m3.
6.3.5 6.3.5.1
TYPICAL DESIGNS
OF
CRYOGENIC STORAGE VESSELS
Vertical Cylindrical Tanks (Figure 6.6)
6.3.5.1.1 General Design These vessels are fabricated with a capacity of up to 13,000 US Gallons (56 t of liquid oxygen). They have an outside diameter of 310 cm with an overall height of 11 m. Internal pressure is approximately 17.24 bar G. Also see the flow diagram (Figure 6.7). 6.3.5.1.2 Materials They may be vacuum insulated with either expanded perlite or superinsulation. The latter form of insulation results in a more compact configuration, and is somewhat lighter in weight. The inner tank containing the cryogen is made either from stainless steel Type 304 or a 9%
FIGURE 6.6 Vertical cryogenic bulk vessel. (Courtesy of Taylor Wharton, 2005. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
Safety valve ASME Safety relief selector SV-1
Rupture Rupture disc disc Dual safety device (optional) V-12 V-2 V-11
PBC-1
CN-1
SV-2
V-4 PCV-1
CV-2 V-14 SV-3 S-1
V-5 V-3
CV-1 CV-2 LI-1 PBC-1 PCV-2 PCV-1 PCV-2 PI-1 R-2 CV-1 VR-1 R-1 CN-4 R-2 S-1 SV-1 SV-2 V-8 SV-3 V-1 PI-1 V-2 V-3 V-4 LI-1 V-9 V-5 V-6 V-8 V-9 V-10 V-10 V-11 V-13 Vaporizer (optional) V-12 CN-5 V-13 V-14 V-6 CN-1 CN-2 CN-3 CN-3 CN-4 CN-5 R-1
VR-1
V-1 CN-2
Legend Economizer check valve Fill check valve Liquid level gauge Pressure-building coil Pressure-building regulator Economizer regulator Pressure gauge, inner vessel Vacuum gauge tube Rupture disc Outer vessel relief device Strainer Safety valve (ASME) Fill relief valve Pressure-building relief valve Bottom fill valve Top fill valve Pressure-building inlet valve Full trycock valve Vacuum gauge valve Evacuation valve Vapor phase isolation valve Equalization valve Liquid phase isolation valve Pressure-building isolation valve Vapor vent valve Vaporizer inlet valve Hose drain valve Fill connection Liquid withdrawal connection Liquid withdrawal connection Vapor connection Auxiliary product withdrawal connection
FIGURE 6.7 Flow diagram of vertical cryogenic storage vessel. (Courtesy of Taylor Wharton, 2005. With permission.)
nickel steel alloy. The inner container must also follow the most recent standards of the ASME Code, Section VIII with full radiographing inspection of welds. If necessary it must also follow the British Standards Institute, Standard 1500 or 1515. In general, these units are designed, fabricated, tested, inspected, and shipped to the site for erection on a prepared slab foundation. Ancillary equipment for these storage tanks includes such items as: vacuum-insulated pipe connections, ambient air vaporizers, and a gas-pressure control assembly connected to a customer distribution system. 6.3.5.2 Horizontal Cylindrical Storage Tanks (Figure 6.8 and Figure 6.9) 6.3.5.2.1 Design and Materials These units are generally an in-house storage system, close to the producing plant and are built up to a capacity of 222,000 L (250 t liquid oxygen). They may be made of either stainless steel Type 304 or a 9% nickel steel alloy. The use of either 5% nickel steel or aluminum is not generally recommended, because the tanks are usually designed for an internal pressure of 1.72 bar G and higher. They are usually vacuum insulated with expanded perlite, thus reducing the evaporated loss to less than 0.5% per day. The fabrication of the inner vessel must follow the most recent standards of the ASME Code, Section VIII, with complete radiographic inspection of welds, and thermal treatment for stress annealing. The outer tank is fabricated with carbon steel, usually 285 Grade C. The completed unit is designed, fabricated, fully tested, inspected, and shipped to the site for erection on the prepared concrete saddles. It is not unusual to find multiple horizontal tanks piped together by a manifold for reasons of economy. Flow diagram per vessel is shown in Figure 6.9.
ß 2006 by Taylor & Francis Group, LLC.
FIGURE 6.8 Horizontal cryogenic storage vessel. (Courtesy of Industrias Brasileira de Gases, 2006. With permission.)
FIGURE 6.9 Instrumentation for a horizontal vessel. (Courtesy of Industrias Brasileira de Gases, 2006. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
6.3.5.2.2 Mechanical Stresses In this design, the engineer should keep in mind that the vessel must withstand the weight of the material of the inner and outer containers, as well as the weight of the contained cryogen. It must also support the bending stresses as the result of the beam-bending action when supported between the saddles of the foundation. Moreover, during the cooling down period when the cryogen is being introduced, there is a possibility of additional thermal stresses, depending on the rate of cryogenic liquid introduction. In regard to the outer tank, the designer must also take into consideration possible shell failure from elastic instability (collapsing or buckling), rather than any excessive stress. The ASME Code covers this type of failure from elastic instability. Charts are given for the design of cylinders and spheres, which may be subject to external pressure. Collapsing or buckling of the outer casing may be due to the great subcooling of the cryogen within the internal vessel. 6.3.5.3 Spherical Tanks 6.3.5.3.1 General Design In the capacity range of 450,000–900,000 L (500 to 1000 t liquid oxygen), the spherical design is possibly the most economical choice. Its design offers the maximum volume for a minimum external surface, in other words, the lowest cost for material for its volume. Each segment of the sphere can be cut, shaped, and the edges beveled for welding under controlled shop conditions, but subassemblies have to be field welded in whatever atmospheric conditions exist at the site during erection. 6.3.5.3.2 Materials and Codes Materials to be used are stainless steel Type 304, and possibly 9% nickel steel, which has a greater resistance to impact at low temperatures than 8% or lower nickel steel. Because of the higher pressures involved, it is not recommended to use a lower content of nickel in any steel. The same code requirements of ASME Section VIII including radiography of welds and stress annealing expressed for pressurized tanks also apply. With regard to stress annealing, the following procedure has been followed: A heavy layer of insulation is applied on the external surface of the inner spherical vessel, with the exception of a small opening. Through this orifice, a propane–butane burner is inserted to supply the necessary heat for stress annealing at the required temperature, in accordance with ASME conditions. Once the code requirements are fulfilled, the vessel is allowed to cool down, the external insulation is peeled off, and the external shell is constructed following the ASME requirements for elastic instability, as mentioned previously for cryogenic double wall tanks. The space between the inner and the outer shells is generally filled with expanded perlite, and is evacuated to a high vacuum. Whereas spherical tanks have proven popular in the past, they have been supplanted to a great extent by flat bottom double wall tanks that are more economical to erect. The latter also offer a smaller loss in boil-off per day. They are limited, however, to applications requiring low pressures. There is an exception to the latter premise as noted in Section 6.1.1 (1) whenever a high pressure and instantaneous supply of a liquid cryogen has to be injected into a pipeline to maintain a process in operation awaiting the start-up of liquid pumps. The normal capacity of such spheres is in the vicinity of 50 liquid tonnes or to maintain the process in operation for 30 min (see Section 6.4.9, Figure 6.10).
ß 2006 by Taylor & Francis Group, LLC.
Inlet filter
Air compressor
D.C.A. Absorbers A
Waterwash tower
Air separation
B Expander Duct
HP LOX storage Tic Fic Ts Tell
LES
FC TC
TC FC FC
Consumer
B Steam fired waterbath LOX vaporizer
LP LOX storage
A HP LOX pumps
Fic
Oxygen compression system
FIGURE 6.10 Emergency spherical vessel system. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
6.3.5.4
Flat Bottom Tanks (Figure 6.11)
6.3.5.4.1 General For large cryogenic storage tanks with a capacity ranging from 800 to 5000 t (liquid oxygen), the advantages of the double wall flat bottom tank are self-evident. The metallic segments for the vertical walls are very easy to roll in the factory, and are easier to weld than spherical segments in the field. 6.3.5.4.2 Design Flat bottom tanks are designed for pressures near atmospheric in the range of 3–34 kPa G. The bottom is not subject to any important stress, except for compression from the contained cryogen and is made from relatively thin material. It consists of a welded metallic plate, which serves as a vapor barrier. On the bottom surface, underneath this plate, there may be applied a layer or two of load-bearing cellular-glass blocks acting as insulation. The inner vertical wall of the cryogenic vessel is, as a rule, anchored to the foundation to hold down the inner shell against any possible lifting forces due to rising internal pressure. Applied constraining devices, however, must allow any displacement of the vertical shell due to thermal variations. 6.3.5.4.3 Materials The material used for the inner shell of flat bottom tanks has been stainless steel, Type 304 originally, but because of the very low pressures involved, this has been replaced by aluminum, then by 9% nickel steel, which in turn has been displaced in some cases by 5% nickel steel. However, the project engineer should be warned. The use of 9% nickel steel has a greater resistance to impact than 8% nickel steel; and that if one considers anything lower than 5% nickel, the project engineer should be advised that around 3.5% nickel, the steel alloy may suffer from embrittlement due to very low temperatures involved, and its use should not be permitted. Flat bottom tanks up to 2500 t of liquid oxygen have been built with 5% nickel steel, but not without the permission of the ultimate user. Its use is one of the balancing
ß 2006 by Taylor & Francis Group, LLC.
13.79 kPa pressure Liquid level
Volume:
2006 m3 530,000 US Gallons
17.20 m
Tank diameter: 12.19 m
Insulation: glass bricks
0.67 m
Pile cap
Piles
3.0 m
LOX pump
1.22 m (a)
Grade level
FIGURE 6.11 (a) Typical design for a large flat bottom Tank. (b) Construction of cryogenic LOX and LIN storage tanks. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
economics with long-term safety. Between the inner and outer shells the space is filled with expanded perlite as insulation.
6.3.6 6.3.6.1
CRYOGENIC LIQUID DELIVERY SYSTEMS General
Once large industrial liquid-producing plants became an economic reality around the end of World War II, the cryogenic industry attacked the next problem, namely, a viable delivery system of the liquid products from the producing plant to the point of use. The transportation of cryogens includes a variety of methods and vessels, ranging from small portable Dewar flasks to large trucks and railroad tank cars carrying liquid oxygen, nitrogen, argon, helium, and even hydrogen across the continent. 6.3.6.2
Small Portable Containers
The capacity of these units may be classified as follows: 1. Very small vessels with a capacity from 1 to 50 L are essentially used for laboratory work. These containers are double wall vacuum insulated, similar to Dewar flasks, which can be handled quite easily. The vessels are at a low pressure. The pressure relief valve is set at around 3.9 kPa G. 2. Heavier containers, which cannot be handled manually, are equipped with casters at the bottom of the outside tank (Figure 6.12). 3. Specially designed portable liquid cylinders are used for small industrial customers, whose consumption of gases is important enough to exceed those described in points (1) and (2). The capacity of these liquid cylinders is in the range of 165 L and operates at
FIGURE 6.12 Large portable liquid helium vessel on casters. (Courtesy of Praxair, Inc., 2006. With permission.)
ß 2006 by Taylor & Francis Group, LLC.
a higher pressure. The relief valve is set at 152.75 kPa G, and the inner vessel safety disk is at 689.5 kPa. These portable, so-called, liquid cylinders can contain up to 148 L (liquid oxygen), at a DOT service pressure of 2014 kPa. Expected evaporation rate is around 1.5% per day for oxygen and 2.2% for nitrogen due to its higher vapor pressure. The inner vessel is generally made from Type 304 stainless steel and the outside container from carbon steel. As a rule they are maneuvered with a wheeled cart designed for that purpose. They can also be manifolded to a common pressure-reducing regulator to serve as an important factory area with high gas consumption. 6.3.6.3 Customer Bulk Stations (Figure 6.6 and Figure 6.7) Whereas these tanks may not belong, strictly speaking, to the category of transportable containers, they are part and parcel of a contractual delivery system. The tanks are rented from the supplier of gases in liquid form and can be removed by the supplier if the contract is terminated. These liquid containers serve hospitals and other large users of gases. Once these portable containers are installed at the customer’s site, they are supplied with liquid product from delivery trucks. The containers are large vertical cylindrical vessels, set on a predesigned concrete foundation pad. They have a capacity of up to 50,000 L, at a pressure of 1724 kPa G (17.24 bar), and an evaporation rate of about 0.22% per day for liquid oxygen. Their overall height is 11 m and outside diameter is 3.1 m. The weight is about 22,600 kg (empty) and 79,000 kg (filled) with liquid oxygen. The vertical tank is equipped with tripods at the bottom, which are anchored to the foundation pad. The top is equipped with shackles for hoisting the tank into place. The inner vessel is made from either Type 304 stainless steel or a 9% nickel alloy steel. The outer tank is from carbon steel, and the insulation is either expanded perlite or superinsulation under a very high vacuum. Fabrication of the inner vessel is by stateof-the-art welding procedures in shop-controlled conditions, including ASME Section VIII applications, stress annealing, and complete weld radiography. The outer carbon steel casing will also include ASME design for preventing elastic instability. Customer stations also include the following ancillary equipments: (a) a low-level alarm to signal a low liquid supply, (b) an ambient air vaporizer that may range from 6 to 135 Nm3=h, depending on the customer’s needs, and (c) a gas flow pressure regulator to control the flow from the supply tank to the client’s inlet pipeline. The system includes a bypass valve, a pressure gauge, and a safety valve (see Figure 6.7). Smaller capacity liquid stations can be completely prepackaged on a skid in the factory, along with an automatic pressure-building circuit and a liquid withdrawal unit. Connections may also be included for high pressure or centrifugal liquid transfer pumps, and a vacuuminsulated liquid pipeline ready to be connected to the client’s inlet pipeline. Smaller liquid stations can be completely prepackaged, instrumented, skid-mounted, as well as shop-tested, and delivered by truck to the client’s site, or by barge to an offshore oil rig. 6.3.6.4 LOX Distribution in a Shop The piping, valves, and accessories should be well-insulated or even vacuum insulated to reduce heat gain to a minimum. 6.3.6.5 Liquid Deliveries by Truck This popular delivery system involves the use of a horizontal cylindrical liquid storage tank, designed and fabricated as described in Section 6.3.5.2, but on a standard truck chassis. Of course, the design and structure of the truck will depend on the final dimensions and the weight of the filled storage tank. The design of the cargo will also involve certain alterations in
ß 2006 by Taylor & Francis Group, LLC.
order to include a cryogenic liquid transfer pump, and any necessary ancillary equipment. Truck deliveries of cryogens generally supply customer stations, or function as a temporary on-site storage facility for special applications, such as a NASA launch site, an oil well site using liquid nitrogen for enhancing oil production, or even for an exchange of products between industrial gas producers. The fabrication of the inner vessel of the container incorporates the use of high-strength materials such as Type 304 stainless steel, or a 9% nickel steel alloy, and conforms to the existing codes of CGA, DOT, and ASME Section VIII. They also require high-technology welding procedures, and very efficient insulation with either expanded perlite or superinsulation, with a high degree of vacuum. The vessel capacity is generally around 16,000 L, or slightly higher, and working pressure is from 2.95 to 19 bar G, to keep product loss low during filling operations. The inner tank is also equipped with transverse baffles in order to minimize sloshing during a long journey on a highway. The tank should also include a high capacity pressure-building coil to facilitate the rapid fill of the receiving tank at the other end of the journey.
6.4 CRYOGENIC PUMPS 6.4.1 BACKGROUND One of the major contributions to the liquefaction of gases has been the development and application of the cryogenic pump that could operate at temperatures below 170 K. It was first applied to small high-pressure liquid-producing plants for the direct conversion of liquid oxygen to a high-pressure gaseous product for cylinder filling. This concept immediately eliminated the old and rather primitive water-lubricated reciprocating compressor, which required weekly maintenance, and was the bane of all operating departments in the compressed gas industry. The original cryogenic pump, introduced in the 1940s, was a positive displacement pump involving a horizontal stainless steel piston housed in a bronze casing, inserted into the cold box casing near the main condenser. The cold box insulation, either rock wool or glass wool, depending on the size of the plant, also insulated the pump. The pump was located close to the casing with a bolted aperture for easy maintenance. The electric motor drive, of course, was located outside of the casing. As good as the original concept was, stoppages due to freeze-ups from entrained solid carbon dioxide particles were not infrequent.
6.4.2
VARIETY
OF
APPLICATIONS
The original small positive displacement pump, whose technology is still used today, has been joined by a whole series of modern centrifugal pumps (Figure 6.13) ranging in application from the simple transfer of cryogenic liquids at low or medium pressures to larger highpressure positive displacement pumps operating at pressures in excess of 200 bar. These applications include: 1. The transfer of liquid products within the cryogenic process cycle itself, and is part of the cryogenic process cycle for process improvement 2. The transfer of liquid products to outside insulated storage tanks for back-up purposes 3. The transfer of liquid products from storage tanks to insulated cryogenic trucks or railroad tank cars for outside distribution 4. The transfer of liquid products from trucks or tank cars to back-up storage tanks 5. Compression pumps working in concert with specially designed vaporizers to convert cryogenic liquid products into high-pressure gas for distribution lines to various points of consumption
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FIGURE 6.13 A cryogenic single-stage recirculating pump as the pressure increases, the number of stages also increases. (Courtesy of Lawrence Pumps, Inc., 2005. With permission.)
6. High-pressure positive displacement pumps operating over 200 bar that are used principally with liquid nitrogen for enhanced oil recovery, but may be considered for oxygen in very high-pressure services Each of these applications may require a specialized design depending on the flow and pressure requirements. It may be said, however, that cryogenic pumps are generally divided into (a) transfer pumps for high liquid flows, and low or intermediate pressures, and (b) compression pumps (including both centrifugal and positive displacement pumps) for the conversion of cryogenic liquids into a vapor phase for gaseous pipeline distribution and direct consumption.
6.4.3 MATERIALS Cryogenic pumps follow the same basic design as pumps manufactured for higher temperatures. Reliable operation below 169.5 K, however, strictly demands use of materials with the proper mechanical and thermal properties at low temperatures. Choice of materials is
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controlled by the following parameters: operational reliability, safety, mechanical configuration, compatibility with oxygen at high pressures, prevention of contamination by gas entrainment or solid particles, and lastly, ease of cleaning and maintenance. Carter3 cited the experiments carried out in Germany by Baur et al.4 in which the following procedures were carried out: . . .
Pumps with aluminum impellers and aluminum casings Pumps with bronze impellers and bronze casings Pumps with stainless steel impellers and stainless steel casings
The operation of the pumps was programed with an axial displacement of the rotor, inducing rubbing between the rotor and the housing. Second, the injection of a foreign material into the pump. In the rubbing tests, two pump geometries were used, both single-stage and with similar tip speeds. The lowest tip speed was 65.85 m=s and the highest tip speed was 77 m=s. This corresponded to a pressure rise of 21.4 and 28.4 bar, respectively. Rubbing was initiated by a special mounting at the thrust bearing, which permitted the rotor to be moved against the housing during the operation. The tests terminated when the torque became high enough to stall the motor. The same tests were repeated with a mixture of liquid and vapor. In the latter tests, both the aluminum combination and stainless steel combination started internal fires, which burned through the housings. There was no ignition or fire with the all-bronze unit. During a test involving an aluminum impeller with a bronze-casing ignition occurred even with only liquid as the medium, but the test with a bronze impeller and an aluminum casing there was no ignition. The use of aluminum, therefore, is not the sole culprit. In the test involving the introduction of a foreign object, the all-aluminum pump caught fire, but the allbronze combination did not indicate any ignition. The preceding tests were carried out with peripheral speeds of about 76 m=s, which include most, if not all, medium pressure and pipeline back-up pumps. In regard to transfer pumps, which operate around 5 bar, a series of tests were also carried out with peripheral tip speeds of 29 to 30 m=s. In these tests, the material combinations used were as follows: An aluminum impeller in a stainless steel volute and an all-aluminum pump. Ignition was attempted by thrusting the impeller toward the housing, and by the introduction of foreign particles into the process. No ignition or fires were observed during the latter tests. One may conclude that up to a discharge pressure of 5 bar, all the above-mentioned material combinations are safe to use against gross axial rubbing conditions. As for medium pressure applications, up to 28 bar, an all-bronze pump will resist gross axial rubs and entrained foreign objects in both liquid- and two-phase operations. However, a pump with an aluminum impeller with a bronze casing is the most sensitive to ignition than all other combinations. There is no doubt that choice of materials and peripheral tip speed are the critical items. A high peripheral tip speed may induce rubbing between the impeller and the casing. The combination of bronze impellers with a bronze casing is the least sensitive to ignition. Carter also reported that radial failures can occur from the rubbing between a rotating shaft and the metal bellows seal, which may be too closely spaced around the shaft. This critical situation may be diminished, however, with the use of a nonrotating bronze sleeve between the rotating shaft and the stationary bellows. In the choice of materials for pumps, one has to bear in mind the application of the pump. For example, if the pump is to handle liquid nitrogen or any other inert fluid, it may be more economical to choose an aluminum impeller with a chrome–nickel steel casing; but if the same
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pump were to have a dual purpose within a few years in the future and were to pump liquid oxygen, then an all-bronze pump is mandatory for safety reasons.
6.4.4 PRESENT DESIGNS As the preceding tests were carried out, all manufacturers of cryogenic pumps have exercised great care in the choice of materials to ensure mechanical reliability at low temperatures and compatibility with oxygen under pressure. Nevertheless, ultimate responsibility rests with the designer, so one must be extremely careful in the preparation of technical specifications for pumps, and in the final acceptance tests at the manufacturer’s shop. It is also important to obtain and study the Compressed Gas Association Pamphlet G-4.15. The design of a cryogenic pump should be as simple as possible to make maintenance easy. Proper selection of seals and bearings is critical. Labyrinth seals with nitrogen injection at one end are commonly specified. Sleeve bearings are used and thrust bearings are lubricated with halocarbon, a synthetic lubricant compatible with oxygen. Impellers are generally made of bronze, and shafts of stainless steel or Monel. Maximum safety casings are made of a tin–bronze alloy. Bronze stops are recommended to prevent the impeller from wandering axially. The impeller should be keyed to the shaft. Electric motor drives, direct gear or belt driven, are normal. For pumping low viscosity fluids such as liquid hydrogen special pumps are required. Their design has to accommodate high differential thermal shrinkages, not only industrially acceptable manufacturing tolerances.
6.4.5 NET POSITIVE SUCTION HEAD6 (FIGURE 6.14) Aside from the importance of a good basic design, the net positive suction head (NPSH) for a cryogenic pump is of greater significance than that for a standard pump operating at a higher temperature. It has to be borne in mind that a cryogenic fluid is always very close to its boiling point. Unless the liquid is kept subcooled, it will vaporize and enter the pump as a two-phase fluid. The expression NPSH may be defined as the difference between the vapor pressure of the liquid and the static head at the suction of the pump. More problems have ensued from this factor than any other single cause. A cryogenic pump cannot be operated on a saturated liquid because in order to establish flow into the pump, there has to be a lower pressure in the pump. If there is a pressure drop in the liquid, the latter will start to boil, and the resultant vapor will enter the pump along with the liquid as a two-phase feed. This condition will create cavitation, and will cause seal or bearing failure. In order to prevent cavitation there must be provided a minimum NPSH at the suction of the pump. According to Martin7, the NPSH available is dissimilar to the NPSH required because the former defines the excess of pressure over the liquid’s vapor pressure at the pump suction flange. With some rare exceptions, centrifugal pumps specify that the NPSH available be greater than the NPSH required for avoiding cavitation. This minimum NPSH depends on the design of the pump, and is usually stamped on the nameplate by the manufacturer. A good rule of thumb for transfer pumps is to specify a minimum NPSH of 1.22 m. Another good rule to adopt is to make sure that all incoming piping beginning from the source of the cryogenic liquid is of ample diameter, minimum length, straight flowing as much as possible, with a minimum of bends and valves. Of course, all the piping should be very well insulated for continuous service. Moreover, the piping on the discharge side should not be undersized. One should allow a maximum velocity of 0.61 m=s, and have a low-pressure drop. In order to stabilize the flow at
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Inducer
FIGURE 6.14 Schematic of an inducer in a pump. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
the entrance port of the pump it is prudent to specify the use of a bronze inducer, which will serve to streamline the fluid flow at the entrance port. This piece of equipment is a small axial type of impeller similar to a conveyor screw, which is placed close to, and on the same shaft as the main pump impeller. The inducer has a very low NPSH, and generates just enough head to provide a suction pressure for the principal impeller. When vaporization occurs at the inlet, the bubbles so formed are small and collapse over the entire length of the inducer. Therefore, when a shock occurs, the inducer keeps operating with no vibration. If no inducer is used the vaporized bubbles implode along a very small portion of the impeller vane because the pressure rise is sudden and steep. The efficiency of the inducer is lower than that of the main impeller, but the effect on the overall operation in terms of energy is minimal. Problems at the pump may also be due to a long and complex piping between the storage tank and the cold box. The liquid cryogen leaving the cold box is generally subcooled at about 1 K or more below its boiling point so as to arrive in storage as a completely saturated liquid. If it arrives slightly above its boiling point one may expect problems at the pump. When a cryogenic pump especially in high-pressure service exhibits a malfunction, the local supervisor may be little too quick to point at cavitation resulting from a low NPSH. The mechanical operation should also be checked for any part malfunction.
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Example In a recent project involving a large LOX storage tank, the high-pressure cryogenic pumps were ordered with a minimum NPSH of 8 ft. (2.44 m). As shown in Figure 6.17 the available NPSH is more than adequate including the pressure drops including the liquid level at the tank, vapor pressure at the top of the tank, and the various valves and piping. NSPH evaluation for LOX back-up pumps Calculate maximum=minimum suction pressure
1. Determine tank’s maximum liquid height Tank diameter 12.19 m; tank capacity 2006 m3 (530,000 US gallons) Cross-sectional area [(12.19)2 3.1416=4]H ¼ 2006 m3, so H ¼ 17.2 m From Figure 6.15 maximum liquid above grade 17.2 þ 0.66 þ 3.0 ¼ 20.86 m Deduct pump elevation 1.22 m Net height 19.64 m Max. suction pressure 19.64 m 9.7927 1.141 ¼ 219.447 kPa Add tank vapor pressure 13.790 233.237 kPa Min. liquid above grade 3.00 m þ 0.66 m ¼ 3.66 m1.22 m ¼ 2.44 m Min. suction pressure 2.44 9.7927 1.141 ¼ 27.26 kPa Add tank vapor pressure 13.79 41.05 kPa LOX pump capacity to be specified on a daily plant production rate: Plant daily production 1088 t=d 1 t LOX 231.5 US gal or 251,880 gal=d Pumping rate 251880=24=60 175 gpm Minimum NPSH required ¼ 2.44 m or 8 ft. or (8=2.31) 1.141 3.95 psi System pressure drop: 1–4 in. tank shutoff manual valve ¼ 2 ft. equivalent drop 1–4 in. tank shutoff automatic valve ¼ 4 ft. 1–4 in. pump inlet manual valve ¼ 2 ft. 4–4 in. elbows 74 ells ¼ 28 ft. 1–4 24 in. long flex hose ¼ 10 ft. 30–4 in. pipe ¼ 1 ft. Total pressure drop ¼ 47 ft., say 50 ft. Pressure drop for 4 in. pipe at 180 gpm ¼ 3.53 ft.=100 ft. length Pressure drop for system ¼ (3.5350)=100 ¼ 1.765 ft. or 1.765=2.31 ¼ 0.764 psi 1.141 ¼ 0.871 psi or 6.005 kPa Actual NPSH required ¼ 8 ft. þ 1.765 ft. ¼ 9.765 ft. say 10 ft. or 3 m Note: Industry does not recommend draining a vessel or tank to the bottom. It generally refills vessel when liquid level descends to approximately 10% of its total volume. Conversions: 2.31 ft. of water ¼ 27.72 in. ¼ 1 psi ¼ 6.8948 kPa 1 m of water ¼ 9.7927 kPa Density of liquid oxygen ¼ 1.141 at 90.18 K 1 US Gallon LOX ¼ 9.522 lbs ¼ 4.319 kg 1 cu. ft. LOX ¼ 71.23 lbs ¼ 32.31 kg All piping calculations have employed US sizing.
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Another problem may exist, however, and this involves the temperature of the liquid oxygen in the tank itself, and the temperature of the product when it arrives at the inlet port of the pump. To take care of this problem the product LOX leaving the process is normally subcooled by 1 or 2 K. In the previous problem it was found that the LOX had been subcooled to 88 K, or two degrees below its normal boiling point of 90.18 K, which ruled out the possibility of cavitation. Finally, when the pump arrived at the factory, it was found that one of the parts had been made from a wrong material. As soon as the part was changed to Monel as originally planned, the pump was returned, installed, and operated perfectly.
6.4.6 INLET FILTER SCREEN The use of a strainer at the pump suction is of prime importance. This screen will trap entrained solids such as frozen carbon dioxide, weld chips, small pebbles, etc. These strainers should be examined and cleaned at regularly scheduled periods. Clogged screens may create sufficient pressure drop to start cavitation even in well-installed pumps. Some of the suggested standards of the Commission Permanente Internationale of Paris are as follows: A strainer should be included in all pumps and should be located as close to the inlet flange as possible. The screen size of the filter should be smaller than the design gap between the impeller and the pump casing.
In general the screen openings should be no more than 0.5–1.0 mm maximum. In fact, some air plant suppliers recommend a screen opening less than 0.3 mm. Moreover, the crosssectional flow area should be 1.5 times the cross-sectional area of the installed piping at the entrance port.
6.4.7 INSTALLATION
AND
CONNECTIONS
A common problem with the installation of all pumps, and especially those operating at cryogenic temperatures, involves the ancillary piping around the pump. Poorly designed piping both at the inlet and outlet can create enough strain to unbalance the pump and lead to malfunction, breakdown, and even a fire. In order to overcome the problems of strains due to piping, one should consider the use of flexible metallic connections at both the suction and the discharge sides of the pump. Of critical importance is the sizing of the piping at the suction side. Too large a cross-sectional area will lead to thermal losses due to a low flow velocity, but inadequate diameter will cause a greater friction loss. It is also recommended that pumps have antivibration mountings to prevent excessive shocks from nearby rolling stock and reciprocating compressors. With cryogenic pumps handling high-pressure liquid oxygen, i.e., 30 bar and over, it is recommended that the units including ancillary equipment such as valve bodies be housed behind a concrete barrier for protection of personnel and equipment (see Chapter 12 on safety).
6.4.8 6.4.8.1
TYPICAL PUMP PIPING CALCULATIONS Velocity Constraints of Cryogenic Fluids
Piping calculations for cryogenic fluids involve an appropriate selection of pipe diameter and fluid velocity, and the project engineer must balance piping costs and constraints of the cryogenic fluid in motion. There are very few published articles on the subject, and advisory
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bodies such as Compressed Gas Association offer a graph on permitted velocities based on safety, but not on efficiency. In practice most designers accept a velocity for liquid oxygen in the range of anywhere from 1.8 to 2.4 m=s, but in a straight horizontal length of pipe. In actual application, however, this velocity drops to 0.61 m=s to allow for a fairly high friction factor due to type of valves, elbows, fittings, and changes in direction. It has been stated that at the exit of a pump a cryogenic fluid such as liquid oxygen becomes subcooled, therefore higher velocities may be used, but then how high? Whereas the previous premise is valid, experience has also shown that once the cryogenic fluid has gone well past the pump, energy generated by the friction factor of the piping, valves, and fittings can well cause volatility, which should be avoided at all costs. Example As piping usually involves US Standards for most applications, the following calculations may be easier to understand: Let us assume that the piping run is 300 ft. (91 m) The vertical head is estimated at 100 ft. (30.3 m) The static head ¼ (12 in. 100 ft.)=24.27 in. LOX LOX density ¼ 71.21 lbs=cu. ft. If LOX flow is 45 gpm And 1 US Gallon of LOX ¼ 9.52 lbs Then weight ¼ 45 gpm 9.52 lbs 60 min ¼ 25,704 lbs=h (3.24 kg=s) Volume of liquid ¼ (45 gpm 9.52 lbs)=71.21 lbs=cu. ft. ¼ 6.016 cu. ft.=min (2.84 L=s) Case A If a 1 in. Schedule 1.0 stainless pipe is selected The internal diameter ¼ 1.097 in. Area ¼ 0.945 in.2 ¼ 0.00656 ft.2 Velocity ¼ (6.016 cu. ft.=min)=(0.00656 60) ¼ 15.28 ft.=s ¼ 4.685 m=s The above velocity is much too high. Case B Select a 2 in. Schedule 10 pipe Internal diameter ¼ 2.157 in. Area ¼ 3.65 in.2 ¼ 0.25347 ft.2 Velocity ¼ 3.95 ft.=s or 1.20 m=s Above velocity is still too high. Case C Select a 3 in. Schedule 10 pipe Internal diameter ¼ 3.26 in. Area ¼ 8.35 in.2 ¼ 0.057986 ft.2 Velocity ¼ 1.73 ft.=s or 0.524 m=s The above velocity is acceptable.
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Case D The same calculation for a 4 in. Schedule 10 pipe will indicate a velocity of 1.01 ft.=s or 0.306 m=s. In terms of economy, select a 3 in. Type 304 Schedule 10 stainless steel pipe. In a 3 in. line with a velocity of 1.73 ft.=s, the following pressure drop calculation is in the following order: Reynolds number (NRe) ¼ (50.6 Q r)=(d m) where r is the density and m is the viscosity. Therefore 50.6 45 gpm 71.21 lbs=cu. ft. 3.26 in 0.12 (cent. poise) ¼ 414480 And the friction factor from graph8 is 0.021 and p=100 ¼ (f r Q2)=D5 or [0.0216 0.021 71.21 (45)2]=(3.26)5 or 0.18 psi=100 ft. 3 ¼ 0.54 psi Executing the same calculation with a 2 in. line it is found that the pressure drop will be 4.2 psi (too high). Assume a 3 in. pipe, and a pump discharge pressure to be 50 psig or 344.75 kPa. In specifying a pump for the above conditions, the following may be used: Suction line ¼ 4 in., Schedule 10, SS 304 Discharge line ¼ 3 in., Schedule 10, SS 304 Fluid is liquid oxygen Inlet temperature is 90.15 K Fluid flow is 50 gpm Head is 50 psig Minimum suction head is 0.2 psig NPSH available is 4 ft. (1.22 m) LOX composition is 99.5% oxygen 0.5% argon 6.4.8.2
Pressure Drops due to Piping Components
There is enough published empirical data available from suppliers such as Crane8 relating to the flow resistance of various components, i.e., valves, tees, elbows, bends, etc. It is known as the K factor, the flow factor is known as C factor, and the friction factor relates to the internal roughness of the pipe7. The pressure drop (Dp) in any piping system depends on: 1. The pipe friction ( f ) or ( ft for complete turbulence), which is a function of the inside diameter of the pipe, its fluid velocity, its density, its viscosity, and the internal roughness of the pipe 2. The change in direction of flow path involving bends, elbows, and change in elevation 3. Any obstructions in the flow path such as control valves, orifice plates, and check valves 4. Changes in cross-section and shape of the flow path whether sudden or gradual Once the resistance of the various components is known, they can be summed and converted into total pressure in terms of psi, or kPa or additional lengths of feet or meters in the
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pipeline. Because most of the data is empirical a certain safety margin should be factored into the calculations. Example In Figure 6.15 it has been assumed that the piping follows the previously chosen 3 in. nominal stainless steel Schedule 10 Type 304 pipe in terms of fluid, flow, and physical properties. In addition, the pipe also includes the following fittings: a 3 in. gate valve, a 3 in. elbow with a 6 in. angle valve. All valves are wide open. Length Internal diameter Fluid and flow Vertical head, fT Viscosity, m Density, r Friction factor Reynolds, NRe Mean velocity
300 ft. (91 m) 3.26 in. (8.28 cm) Liquid oxygen at 90.18 K and 45 gpm (2.84 cm=s) 100 ft. (30.3 m) 0.19 71.23 lbs=cu. ft. (1141 kg=m3) 0.018 261,778 1.73 ft.=s (53 cm=s)
Pumping system of 2500 tonne LOX tank total liquid drop 6.706 m Inner tank size 12.19 m diameter 17 m high 2.54 m
Center line of suction
43.2 mm
5.309 m
1.83 m
1.46 m to top of grating
1.6 m
1.27 m
1.37 m
Ground level
1.96 m to flange
Pump suction flange 48.25 mm Ground level
0.915 mm Pump suction
FIGURE 6.15 Example calculations for NPSH. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
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Case A The piping system as described but following a horizontal plane in space9. Gate valve K factor 8 fT or 80.018 ¼ 0.144 Elbow K factor 8fT or 120.018 ¼ 0.216 Angle valve K factor 150fT or 1500.018 ¼ 2.85 Total K factor for fittings ¼ 3.21 Pressure drop for all fittings ¼ [181063.21 71.23(45)2]=(3.26)4 ¼ 0.074 psi (0.51 kPa) K factor for piping ¼ (fT L=D) or (0.01830012 in.)=3.26 or 19.88 Pressure drop for piping ¼ [18106 K r Q2]=(3.26)4 ¼ 0.457 psi (3.151 kPa) Total pressure drop for piping and fittings is 0.074 þ 0.457 ¼ 0.531 psi (3.661 kPa) Case B If the elbow at 175 ft. (53.8 m) is raised at 908 with a 6 ft. (1.83 m) radius so that the pipe section from elbow to angle valve 50 ft. (15.25 m) is vertical. The last 75 ft. (23 m) is once again on the horizontal. All the K factors for the valves and fittings remain as before. The K factor for the piping remains the same, but the change in elevation has to be factored in. Pressure drop for 50 ft. rise ¼ (50 ft. r)=144 in.2 ¼ (50 71.23)=144 ¼ 24.73 psi (170.5 kPa), which should be added to that in Case A or 0.531 þ 24.73 ¼ 25.26 psi (174.17 kPa). Case C The same case as in Case A but the 3 in. gate valve is replaced with a 3 in. globe valve with a K factor of 340 0.018 or 6.12. In this case the resulting K factor becomes a net total (340 – 8) (0.018) or 5.97, and means an extra pressure drop 0.137 psi (0.95 kPa) will be added to Case A. In summary, the project engineer should make an effort to select valves and fittings with the lowest possible K factor, piping with a very smooth internal surface, and avoid as much as possible any rise in elevation. During execution of the piping system strict supervision should also be the key word to avoid additional resistance to fluid flow from poor workmanship.
6.4.9
START-UP
OF
PUMPS
Before starting up any cryogenic pump, it should be cooled slowly and completely down to the temperature of the liquid. One has to make sure that the impeller can operate freely at all
Angle valve
B Elevation 50 ft. 15.25 m
A Get valve
Radius 6 ft. 1.8 m Elevation zero
175 ft. to bend 53.4 m
FIGURE 6.16 Pressure drops due to piping components (i.e., valves, tees, elbows, etc.). (Courtesy of Crane Company, 1986. With permission.)
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times. In some operations a transfer pump is completely immersed in the liquid to be pumped to make sure that, when the start-up signal is activated, the pump will start-up immediately. In this type of operation, the vertical type of transfer pump may have an advantage over the horizontal configuration. In some field applications the cryogenic pump will not start immediately, and may require anywhere from 5 to 30 min to start pumping regardless of what the pump supplier may claim. (Inadequate priming of the pump is the principal reason for unsuccessful starts.) All cryogenic pump installations must be equipped with a valved vent line and a drain at the suction flange of the pump. If the process involving a cryogenic pump imposes a maximum cutoff time of less than 10 s, as may well be the case in the partial oxidation of hydrocarbon feedstocks for synthetic fuels, it is prudent to connect the oxygen distribution line directly to a small pressurized liquid storage tank (see Figure 6.10). In case the pump does not activate immediately, the storage tank will feed a gaseous product into the pipeline for 30 min at least, so that the main process does not stop. During start-up, the pump will begin to rotate on its own accord due to the gas flow through the pump. This phenomenon is caused by the initial and possibly appreciable vaporization of the boiling liquid. As long as this rotation is well below the operating speed of the pump, the overall effect is negligible. It is imperative, however, to make sure that this rotation does not accelerate close to the design speed. Cooldown may be judged as complete when liquid begins to emerge for the vent line. The vent line must be kept open all through the start-up, and shutoff only when liquid begins to appear. The start-up is then complete. If the pump has been idle for a longer period of time, say over a month, it may be prudent to turn on the sealing gas (dry oil-free nitrogen) for several hours before beginning the cooldown operation, so that if any moisture that enters the machine will be removed, thereby avoiding any icing conditions.
6.4.10 HIGH-PRESSURE RADIAL PUMPS (FIGURE 6.17a,b) High-pressure centrifugal cryogenic pumps have had a long history of successful operation since the early days of partial oxidation, in the 1950s, when reliable and safe oxygen compressors were not yet available. These pumps are vertical multistage pumps involving anywhere from 3 to 21 stages depending on the flow and pressure required. They have been used in various services including oxygen, nitrogen, argon, and liquefied natural gas. They can be used in two ways and are as follows: 1. First, as a column pump within the process extracting liquid oxygen from the main condenser, pumping at a high pressure through the primary heat exchangers where the liquid is vaporized in countercurrent heat exchanger with incoming warm air. In this process cycle the primary heat exchangers must be designed to withstand the high product pressure. For a pressure of over 30 bar indicates a higher plant investment. Moreover, the same process demands a higher energy consumption, around 8%, because of the loss of the latent heat of vaporization during the change of phase from liquid to vapor. This concept must be evaluated against the use of a standard low-pressure Claude cycle equipped with an oxygen compressor, which is always more efficient compared to pumping cycle. The Claude cycle, however, may not be found viable for plants where the high cost of an oxygen compressor, with all its ancillary piping and controls, makes its application uneconomical. This has been found to be the case whenever the oxygen plant is tied in with a liquid nitrogen-scrubbing unit for the production of ammonia. In the latter process, the pumping cycle is economically competitive.
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Pump can purge out Insulation fill Connector Nitrogen purge inlet Gasket Fiberglass batting
Barrel drain
Pump can purge in Discharge Perlite fill connector Nitrogen purge exhaust filter 10.2 mm
Grade
Pump mount barrel and insulation 81.3 mm diameter Perlite insulation Insulation Outer barrel 51 mm Pump can 20.3 mm
(a)
Nitrogen sparger End plate (wrap for protection) Concrete base
Per pump length
15.25 mm
FIGURE 6.17 (a) Pump mount barrel. continued
2. Another application of the high-pressure pump is to serve as a back-up, supplying large quantities of product to keep the overall process in operation during a temporary emergency shutdown of the air separation plant. In this case, the pump is tied in with a large liquid product storage vessel at the intake, and linked with a high volume steam vaporizer at the discharge side. The storage tank supplying the pump must be designed in such a manner to deliver enough products for at least 2 d depending on the longest possible shutdown foreseen for the oxygen plant. A 3 d supply is more prudent, especially for a metallurgical operation wherein any shutdown is very costly. Capital cost of liquid storage may be minimized by supplying liquid product from large tanker trucks from nearby merchant plants. The cost of this option is very high, however, both in terms of capital and anxiety.
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LOX from tank
Break out spool Vent, etc. Connector
Add spool for additional connections if required
Vacuum jacketed valves Support
Support
(b)
FIGURE 6.17 (continued) (b) Cryogenic pump connections. (From Kerry, F.G., Inc.)
6.4.11 ULTRAHIGH PRESSURE PUMPS In recent studies for the production of synthetic fuels, the compression of large quantities, up to 2000 t=d of oxygen to pressures of 100 bar, and even to 140 bar, became a problem. Oxygen compressors are already operating at pressures of up to 60 bar, but above this pressure seal losses become very significant, and increase geometrically as the pressure gets higher. In these conditions efficiencies drop markedly. In fact, the projected efficiencies for centrifugal compressors as a function of pressure are given in Table 6.3. An experimental radial machine has been developed to compress oxygen from 60 to 125 bar with a separate casing housing four extra stages. Polytropic efficiency was stated as 66%, but the average was closer to 60%, as projected in Table 6.3. The cost of such an industrial
TABLE 6.3 Projected Efficiencies at Various Pressures Pressure (bar) 40 70 80 100 125
Efficiency (%) 63–65 62–64 61–63 60–62 59–61
Source: Courtesy of F.G. Kerry, Inc., 1977. With permission.
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machine may prove excessive, as it used Monel metal for all piping and valves. A review of a pumping cycle with reciprocating pumps may well worth the effort. Such pumps have already proved their worth in the pumping of nitrogen up to 145 bar for enhanced oil recovery, but with relatively small flows (2–4 gpm). They are rugged, easy to maintain, and can be operated in tandem and at a relatively low capital cost. The design of a multistage centrifugal cryogenic pump may not be described here, because of economic and safety considerations. According to Karassik et al.9, the design parameters for a high-pressure pump have the following relationship: Q2 ¼ Q1 (D2 =D1 ) and
DH2 ¼ DH2 [(D2 )2 =D1 ]
(6:1)
It may involve a pump flow of 300 gpm (19.14 L=s of liquid oxygen), an impeller diameter exceeding 254 mm (10 in.), and a tip speed of 47 m=s. Moreover, the shaft can have a length of around 3 m, and support approximately 24 stages, which poses quite a design as well as an operating problem. The design for shaft bearings as well as for the motor will be a major problem in terms of safety according to Baur et al.4. Added to this is the question of seals. The choice of materials for high-pressure cryogenic pumps follows the same parameters as those for low- and medium-pressure transfer pumps in terms of operational reliability and safety. In the case of vertical multistage pumps, however, extra care must be taken in the design of the shaft and impeller to avoid rubbing against the casing. The choice of metals is most critical. For example, a greater use of softer and more ductile metals such as Monel and cuprous alloys with a minimum content of 40% nickel when in contact with stainless steel is highly recommended. In summary, before concluding that liquid oxygen pumping is the way to go one must keep in mind that liquid oxygen pumping extracts from the cryogenic system the ability to recover the equivalent of the product’s latent heat at its operating conditions, resulting in a fairly large cold end heat load. This loss will be a power penalty of 3%–14% depending on the final product pressure.
6.4.12 AUTOMATION As with pumps for higher temperature service, cryogenic pumps can be computer-controlled from a central distributor system, or put under fully programable logic controls (PLC) and distributed control systems (DCS) for cooldown operation. Proximity probes will indicate any axial movement of the rotor assembly, forestalling radial-bearing failure, and imbalance of rotating members. A resistance temperature detector (RTD) may be used as an automatic shutdown device. Monitoring bearing temperature is necessary on the fan side and especially on the pump side. Pump side-bearing temperature must not be allowed to drop below 24.8 K. Specified minimum and maximum bearing temperatures should be relayed to the central control system and to an alarm and trip.
6.5 6.5.1
CRYOGENIC LIQUID VAPORIZERS10 GENERAL OVERVIEW
When oxygen, nitrogen, or argon are liquefied for back-up storage or transport to a distant point the cryogens must then usually be vaporized for use. A wide variety of equipment is available for vaporization. Selection will depend on the locale, utilities available, and cost. Such equipment falls into two categories:
ß 2006 by Taylor & Francis Group, LLC.
. .
Low duty ambient air vaporizers High duty steam and hot water vaporizers
6.5.2 AMBIENT AIR VAPORIZERS (FIGURE 6.18) These vaporizer heat exchange elements are generally made of extruded aluminum fins called Thermofins. They have a tapered geometry, with a cross-section surface similar to a star. Each of these Thermofins has a central core, which serves as a passage for the cryogen to be vaporized. Several of these fins are attached to form an individual module. The vaporization capacity of each module depends on the number of Thermofins that are in the module. The cryogen is distributed through the central core of the Thermofins in the module, in heat exchange with the ambient air passing around the combined external surface of the module. Depending on the specific design of the various groupings, modules can be rated to vaporize anywhere from 170 to 1340 Nm3=h of gaseous product. Using stainless steel for the internal cores, the cryogen can be pumped and vaporized at a higher gaseous product pressure (up to 98 bars), but of course, at a slightly lower rate of vaporization. 6.5.2.1 Modules Spread Apart A more recent design of the ambient air vaporizers has involved the use of a wide gap between two groupings of the Thermofins in the same module, extending the time necessary to shut down the operation for purposes of defrosting. In this manner, the new design, called Super Gap Thermofins, extends the operating time to be at least 60%. 6.5.2.2 Modules in Alternate Operation In a further development of the same design of the ambient air vaporizer, two modules can be piped in a parallel formation and operated alternately, so that one module can be in operation whereas the parallel module is undergoing defrosting. This design can be operated automatically in a manner similar to reversing exchangers in a standard air separation plant. 6.5.2.3 Modules with Pressurized Air Other variations of the same vaporizer involve the use of a cryoduct fan-boosted ambient air vaporizer with a unit capacity of 2700 Nm3=h, and the use of a defrostable fan ambient vaporizer with two units and an automatic electric defroster for continuous operation.
6.5.3 DIRECT STEAM VAPORIZERS (FIGURE 6.19) Also available are high-performance automatic vaporizers using steam ranging from 0.7 to 10 bar (70–1000 kPa). These vaporizers are of special interest to industries such as steel mills, which have a fair quantity of waste steam available at 3 bar (300 kPa). These units have a vaporization capacity ranging from 700 to 34,000 Nm3=h (oxygen). The main components of these high-performance vaporizers involve: . . .
.
A horizontal steam chamber A removable U-bend tube bundle of 18-BWG-stainless steel tubes A gas discharge temperature controller connected to a pneumatic steam valve and a condensate trap assembly A computerized full ASME VIII Division 1 Code design
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Useful range days before excessive ice clog 6 21 Cold/dry 2 7 (winter) Warm/humid (summer)
10
O2/N2 flow (SCFH)
7
70,000
5 3
30,000
Superg
ap (S-G
2
)
Sta nd
ard
1 10,000 7 5
5,000
0.1
0.2
60% Gain
0.3 0.5 0.7 1.0 2 3 5 7 10 14 21 Days continuous operation
28
FIGURE 6.18 Super gap vaporizers. (Courtesy of Thermax Corporation, 2000. With permission.)
Operation is fairly simple. Liquefied gas enters the tube and is vaporized and superheated by the steam condensing in the steam chamber. The warm gas leaving the unit activates the controller, which in turn throttles the steam valve. The final condensate trap drains the spent steam from the chest when holding the steam under pressure within the chest. At lowtemperature settings, or low throughputs, a vacuum breaker on the chest admits air in it, cutting heat to the liquefied gas, and preventing condensate buildup in the chest. The attached diagram illustrates the equipment involved. 6.5.3.1
Vaporization with Steam-Heated Water
Another variation involves the use of a steam-heated hot water bath (Figure 6.19). This application is useful if more than one cryogen needs to be vaporized at different capacities
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Steam vaporizer supplied by thermax
Steam in
Pneumatic steam control valve
Steam strainer
Steam shell relief set Valve bypass (optional)
SI H
Options
Gas discharge N2 N4
Vacuum breaker
RE Gas discharge
Temp. control instr. air 1.7 Barg. trim FPT
Steam valve a/o trap block and bypass valves prop and reset temperature controller low temperature shutdown
N3
Weatherproof insulating jacket
Drip trap connector 13mm FFT
RE Liquid in FE
N1 Main liquid in
CI
Other ratings Main trap Ethane: same as connector propane
L
Bolt hole 4 PLCs
Leave (5 + 50 cm) clear to pull bundle
Pressure drop S style 1.4 BARD at 13.8 BARG/SH style 2 BARG at 13.8 BARG For 75% flow use 1/2 of these values Other values call thermaz for bulletin PC-13
FIGURE 6.19 Steam vaporizer. (Courtesy of Thermax Corporation, 2000. With permission.)
simultaneously. Two or more separate coils, each carrying a different cryogen, are submerged in the same bath heated by steam or electrically, depending on the local cost of utilities.
6.5.4 EMERGENCY VAPORIZATION
OF
PRODUCTS
Another application is also very useful. For example, in an emergency shutdown of the local air separation plant, a product flow from a small high-pressure storage tank can be released and vaporized immediately in order to maintain product flow, whereas the emergency highpressure product pumps are activated to supply product from the larger storage tank. In such a situation, the small auxiliary high-pressure storage tank should have a capacity of around 30 min supply of product.
REFERENCES 1. Springman, H. 1978. The liquefaction of oxygen, nitrogen and argon. Linde reports on science and technology, No. 28, pp. 23–30. 2. UBC Uniform Building Code Division IV Earthquake Design. 3. Carter, T.A. 1971. Oxygen Compressors and Pumps Symposium, November 9–11, 1971. Compressed Gas Association Inc. pp. 57–60. 4. Baur, H. et al. Fire tests on centrifugal pumps for liquid oxygen. Kaltechnik, Klimatisienung No. 4, 1970, and No. 4, 1971. 5. Compressed Gas Association, Installation guide for horizontal. Stationary electric-motor driven liquid oxygen pumps, 1st ed., CGA G-4.7, 2006.
ß 2006 by Taylor & Francis Group, LLC.
6. Grohmann, M. 1979. Extend pump applications with inducers. Hydrocarb Process (December) 121–124. 7. Martin, G.R. 1996. Pumps and NPSH: Avoid problems and improve reliability. Hydrocarb Process (May) 63–65. 8. Flow of fluids through valves, fittings, and pipe, 25th ed., Technical Paper No. 410 and 410 M, Crane Publication, 1991. 9. Karassik, I.J. et al. 1976. Pump handbook, 2–32. New York: McGraw-Hill. 10. Bernert, R. Thermax Incorporated, 3 Pleasant Street, South Dartmouth, MA 02748.
ß 2006 by Taylor & Francis Group, LLC.
7
Insulation
7.1 GENERAL In cryogenic systems for the liquefaction, separation, and purification of gases, one has to deal with temperatures of around 112 K for liquid methane, and as low as 4.2 K for liquid helium. The design engineer, therefore, has to provide adequate insulation for the processing and ancillary equipment subjected to these very low temperatures, in order to reduce the heat gain to a practical minimum. Any heat gain from the outside ambient atmosphere means a corresponding loss in refrigeration applied to the process cycle. Since the cost of externally applied refrigeration, energy for compressors, mechanical refrigeration units, far outweighs the cost of insulation, it is therefore prudent to spend the necessary time to study the requirements for insulation, instead of leaving it in the hands of others or of the cheapest insulation supplier. The lower the processing temperatures, the greater the need for properly specified insulation, in terms of type, quantity, quality, and density. For industrial purposes any temperature lower than 120 K may be considered as in the field of cryogenic processing. Cryogenic process systems may involve any or all of the following materials and equipment operating at very low temperatures: 1. Processing equipment, such as heat exchangers, distillation columns, and piping, enclosed inside steel casings commonly referred to as the cold box 2. Externally located process or transfer piping 3. Storage vessels for cryogens (cryogenic liquids) All these elements require properly specified insulation in order to reduce heat gain (heat flow from the outside ambient conditions into cold equipment), and thus maintain a low-energy consumption for the process. Moreover, each element may require a separate study because of the size, configuration, and duty of the specific elements.
7.1.1 THEORETICAL CONSIDERATIONS Although the term heat gain (W=m2) is synonymous to heat flux and heat transfer, the former should be used for cryogenic systems because it implies that heat is entering the system. The basic principles still involve the same concepts as heat transfer. Therefore, a quick review of these principles and their terminology may be useful. This chapter is not intended for design purposes, but for understanding and evaluating various insulations available in the present market, and their proper selection.
7.1.2 INSULATIONS: GENERAL The thermal conductivity of any insulation is based on interaction of any or all of the following parameters: 1. Thermal conductivity of the solid components forming the insulation material or structure 2. Thermal conductivity of air, or of any residual gas within the insulation
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3. Heat transfer by natural convection within the cells or pores of the material 4. Heat transfer by radiation within the structure of the material
7.1.3 VACUUM INSULATION (RADIATION) Heat transfer with no material carrier Vacuum insulation is almost synonymous with the term Dewar flask named after James Dewar, who in 1892 not only conceived the idea of vacuum insulation, but also built the first double wall vacuum-insulated flask. However, the theory behind the concept is a little more involved than the original basic product. This is especially true in the use of multilayer insulation (MLI), which involves a large number of reflecting surfaces. The Stefan–Boltzmann law deals with the quantity of heat transferred by radiation, Q, from a blackbody that absorbs all radiant energy, which strikes it. It is proportional to the fourth power of the body’s absolute temperature. Since a blackbody absorbs and emits the maximum amount of energy at a given temperature, it is necessary to introduce an emissivity factor Fe to estimate the radiation of an actual body. The emissivity is the ratio of the actual radiation to the radiation of an analogous blackbody. To calculate heat gain from radiation, the modified Stefan–Boltzmann equation may be stated as follows1: Qr ¼ sFe F12 A1 (T24 T14 )
(7:1)
where Qr is the radiant heat transfer rate per unit area, s is the Stefan–Boltzmann constant (56.69 nW=m2=K4 (n ¼ nano ¼ 109)), Fe is the emissivity factor, F1–2 is the configuration factor, A1 is the area of colder surface (m2), and T is the absolute temperature (K). For standard cryogenic double wall vacuum-insulated vessels F1–2 ¼ 1, where subscript 1 refers to the inner vessel, and subscript 2 refers to the outer vessel. The emissivity factor for diffuse radiation for concentric spheres or cylinders is given by 1=Fe ¼ [1=e1 þ (A1=A2 )(1=e2 1)]
(7:2)
where e1 is the emissivity of the colder surface, A1 is the area of the colder surface, e2 is the emissivity of the warmer surface, and A2 is the area of the warmer surface. If a thermally isolated floating highly reflective shield is placed between the warm and cold surfaces, then the emissivity factor becomes as follows: 1=Fe ¼ (1=e1 þ 1=es 1) þ (N 1)(2=es 1) þ (1=e1 þ 1=es 1)
(7:3)
This equation shows that heat gain can be reduced tremendously. For example, in the basic equation with no radiation and no shields, if e1 ¼ e2 ¼ 0.80, then Fe ¼ 0.6667, but if 10, highly reflective shields each with an emissivity of 0.05 were interposed, then the overall emissivity becomes 0.00255. In other words, radiation heat gain is reduced by a factor of 0.00383. The advantage of using reflecting shields with any insulation becomes obvious as will be shown in MLIs (Figure 7.1). In a vacuum insulation the principal heat gain by radiation is also supplemented from a smaller heat gain by free molecular conduction. The residual molecules remaining in the evacuated space carry energy from the outer warmer surface to the colder inner surface without transferring energy to other molecules. They never remain on either surface long
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T2
! T1
T1
Warm surface
Cold surface
T1
FIGURE 7.1 Energy transportation from free molecular movement. (Courtesy of F.G. Kerry, Inc., 2006. With permission.)
enough, however, to establish a thermal equilibrium. This inefficiency, therefore, is expressed by the accommodation coefficient a, which is a ¼ actual energy transferred=maximum possible energy transfer The accommodation coefficient a is based on a specific gas-surface combination, and the surface temperature at the warm end and at that on the cold side. For example, a1 ¼ (T21 T11 )=(T21 T1 ), at the cold surface a2 ¼ (T21 T11 )=(T2 T11 ), at the warm surface
(7:4)
T2 T1 ¼ (1=a1 þ 1=a2 1)(T21 T11 ) ¼ (T21 T11 )=Fa
(7:5)
and
Fa is the accommodation factor that has the same form as the emissivity factor, and can be calculated in the same manner 1=Fa ¼ 1=a1 þ A1 =A2 (1=a2 1)
(7:6)
Approximate typical accommodation coefficients are given in Table 7.1.
TABLE 7.1 Accommodation Coefficients for Air, Hydrogen, and Helium Temperature (K) 300 77.7 20
Air
Hydrogen
Helium
0.8–0.9 1.0 1.0
0.29 0.53 0.97
0.29 0.42 0.59
Source: From Timmerhaus, K. and Flynn, T. in Cryogenic Engineering, Springer Science & Business Media, New York, 1989. With permission.
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To determine the total heat gain of a typical double wall vacuum-insulated vessel, the following factors must be considered: 1. The emissivity factor, 1=Fe ¼ [1=e1 þ (A1 =A2 )(1=e2 1)]
(7:2)
2. The radiant heat transfer, Qr ¼ sFe F12 A1 (T24 T14 )
(7:1)
3. The accommodation coefficient factor, 1=Fa ¼ [1=a1 þ A1=A2 (1=a2 1)]
(7:6)
4. The heat gain from free molecular conductivity heat gain, or Qc (for concentric cylinders and spheres). The energy supplied by the free residual molecules is given by De ¼ (cv þ (1=2)R)(T21 T11 )
(7:7)
where R is the specific gas constant (R ¼ Ru=M), Ru is the universal gas constant (8.31441 kJ=kmol K), M is the molecular weight of the gas, cv is the specific heat of gas [R=(g1) for an ideal gas], and g is the specific heat ratio, cp=cv (g ¼ k at times). Eliminating effective molecule temperatures, and from the kinetic theory of gases, the mass flow rate of molecules per unit of surface is m=A ¼ p(T=2pgc R)1=2
(7:8)
where T is the gauge temperature of gas (K), and gc is the conversion factor in Newton’s second law of motion, which states that acceleration of an object is directly proportional to the resultant forces acting on it and inversely proportional to its mass, or SF ¼ ma. This is a vector expression, hence proportional to three component equations. The equation is valid only when the speed of the particle is much less than the speed of light. In SI units, mass ¼ kg, acceleration ¼ m=s2, force ¼ N ¼ kg m=s2, and N ¼ 105 dyn (9.80665 m=s2). The energy transfer rate by molecular conduction is therefore Q=A ¼ (m=A)De
(7:9)
or Q ¼ GpA1 (T2 T1 ) and G ¼ [(g þ 1)=(g 1)][(gc R=8pT)1=2 Fa
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(7:10)
Free molecular conductivity to occur, however, the mean free path (l) of the gas molecules must be larger than the spacing between the two surfaces. To determine this l ¼ (m=p)(pRT=2gc )1=2
(7:11)
