team 32 - overall team report

CN 4120: DESIGN II Project

PRODUCTION OF HYDROGEN VIA SYNGAS ROUTE

Overall Team Design Report

TEAM 32:

Lim Yueh Yang U046787U (Steam Methane Reformer)

Ng Su Peng U046929L (Furnace)

Ong Song Kun U046829M (High Temp. Shift Reactor)

Tham Zhi Yong, Andrew U046754W (Low Temp. Shift Reactor)

Zhang Zihong (Leader) U046816H (Pressure Swing Adsorption)

Sin Yew Leong U046835M (Heat Exchanger Network)

Heng Chee Hua U046793U (Cooling Tower)

This report is submitted in partial fulfillment of the requirements for the Degree of Bachelor of Engineering (Chemical)

Department of Chemical & Biomolecular Engineering National University of Singapore

2007/2008

CN 4120: Design II Team 32 Executive Summary

Production of Hydrogen via Syngas Route

EXECUTIVE SUMMARY

Hydrogen is vital for daily operations in refineries worldwide due to its primary usage

in hydrotreaters, as environmental regulations on sulphur emissions are strictly enforced.

Furthermore, secondary units, such as hydrocrackers are constructed to boost the margins of

refineries through upgrading of middle distillates (kerosene and diesel), which requires the

hydrogen. Thus this report aims to develop a preliminary design for a hydrogen plant (1.25e9

m3 (STP/year)), whose operation is based on the syngas route that involves the coupling of

steam-methane-reforming with low temperature and high temperature shift reactions. Further

assumptions such as siting the plant in Singapore and an 8000 h/year operation time were

also considered in the production of hydrogen with at least 99.9% in product purity.

The preliminary design consists of seven main units – furnace, steam-methane-

reformer (SMR), high-temperature-shift-reactor (HTS), low-temperature-shift-reactor (LTS),

pressure swing adsorber (PSA), heat exchanger network (HEN) and cooling tower.

The main purpose of the furnace was to supply heat to the feed mixture of steam and natural

gas so that the endothermic reaction can proceed in the SMR. Optimal design indicated the

requirement of a four-chamber single-tube-pass side-fired heater which comprised of a

radiant section (33m x 27m x 13m), a convection section (7m x 7m x 1m) and a stack

(diameter = 4m, height = 8m). A thermal efficiency of 94% was achieved. The refractory

walls comprised of firebricks with silicon carbide linings.

450 SMR tubes (material = HK-40) were housed in the radiant section of the furnace

where 6.0e8 kJ/h was supplied via the combustion of the tail-gas directed from PSA and

excess air. A steam-to-carbon ratio of 3:1 was stipulated, which would also minimize coking.

Prior to entering SMR, the feed was preheated to an inlet temperature of 539oC. In the

presence of Ni/Mg-Al2O4, methane would react with steam to produce an effluent that

contained primarily carbon monoxide and hydrogen. A methane conversion of 80.1% was

attained with an exit temperature of 852oC. The total cost of the SMR tubes and furnace were

estimated to be US$730,000 and US$37 million respectively.

The SMR effluent was subsequently cooled before entering HTS at a temperature of 354oC,



with make-up steam to achieve a steam-to-carbon ratio of 5:1. The HTS vessel (diameter =

3.46 m, height = 12.11 m, material = ASTM A387) served to increase the hydrogen yield

through the oxidation of carbon monoxide to carbon dioxide, in the presence of chromium

promoted iron oxide. The stream composition of carbon monoxide was subsequently reduced

from 13.3% to 3%. The calculated bare module cost was approximately US$3.6 million.

As the oxidation process was slightly exothermic, a lower temperature operation

would favour a higher conversion, thus justifying further process cooling to 220oC prior to

entering LTS. Through the optimal design of the LTS vessel (diameter = 3.31 m, height =

5.07 m, material = ASTM A387) with copper-zinc oxide catalyst supported on alumina, the

carbon monoxide level was further lowered to 0.5% at the LTS exit. To prevent poisoning of

PSA catalyst downstream by condensate, a knock-out drum and a bed of silica gel was

installed after LTS, prior to the entry into PSA. The bare-module cost of LTS was estimated

at US$1.4 million.

Further cooling of the LTS effluent to 50oC was effected before entering into the

knock-out drum and subsequently into PSA. The composition of the PSA feed was roughly

75% hydrogen and 18% carbon dioxide. A Polybed system of 8 columns (diameter = 3 m,

height = 8.5 m, material = SS clad), operating between 1 and 25 bar at 50oC, was adopted for

the purification of hydrogen with the use of activated carbon and zeolite 5A at a ratio of 5:1.

A hydrogen recovery of 85% with a product purity of 99.9% was subsequently achieved. The

estimated bare module cost of PSA was $25 million.

Extensive heat integration was performed for maximum energy recovery in this

design. Only cold utilities, such as high pressure steam and cooling water were needed, as the

furnace had fulfilled all the heating requirements of the plant. This resulted in the presence of

a utility pinch, which requires the adoption of pinch analysis. 3 networks each satisfying the

maximum energy recovery criterion was designed using HX-Net. The selected network was

chosen based on the lowest total annual cost and operational considerations, attaining

100.6 % of the total cost target. The chosen TEMA configuration of the heat exchangers was

that of AES shell and tube exchangers (split-ring floating head). For thermal design, the heat



exchanger chosen possessed a heat transfer area of 519.2 m2 (calculated by HX-Net) with 456

SS tubes of length 4.88 m while the shell was fabricated from carbon steel. The cost of this

heat exchanger was about US$526,000.

A cooling tower was designed based on an induced draft counter-flow configuration.

A filled height of 6.0 m was essential for the rejection of heat into the atmosphere via both

evaporation and sensible means. Replenishment of water (188 m3/h) was required for

continuous operation due to evaporative losses. Construction costs were estimated at

US$941,000. The lifespan of all catalysts was assumed to be 3 years and their cost amounted

to roughly US$2.2 million/year. Based on a discounted cash flow rate of return of 10% and a

payback period of 15 years (inclusive of 2 years of construction), the selling price of

hydrogen calculated was US$2.43 / kg (STP), which was less than US$2.70 / kg (STP)1.

Therefore, we recommend the construction of the plant due to the profitability of the product.

Safety is paramount, thus a HAZOP worksheet was generated to identify potential

hazards due to possible deviations in both SMR and furnace operations. Recommended

safeguards and actions were also highlighted. A summary of occupational safety and health,

environmental impact assessment and plant layout was further discussed in this report. Lastly,

an implementation of process controls and instrumentations was performed on both SMR and

furnace. A piping and instrumentation diagram (P&ID) was subsequently developed with

further discussions centering on the various reflected control strategies.



REFERENCE

1. Hydrogen and Clean Fuels: Systems Studies. Retrieved on April 17, 2008 from National

Energy Technology Laboratory Web site:

http://www.netl.doe.gov/technologies/hydrogen_clean_fuels/systems_studies.html

ACKNOWLEDGEMENTS

This section dedicates acknowledgements to all who have helped our team by offering

their valuable advice. In particular, we would like to express our heart-felt gratitude to our

professors, Prof Karimi, Prof Rangaiah, Prof Farooq, A/P Kawi, A/P M.P. Srinivasan, A/P R.

Srinivasan, A/P Krishnaswamy, A/P Hidajat and A/P Borgna, for their valuable insights.

Last but not least, this work would not have been possible without the coordination

and teamwork in Team 32 (CN4120 – Sem 2 of AY2007/08), hence we would like to thank

all of them for their assistance and understanding.

CN 4120: Design II Team 32 Process Flow Diagram (P.F.D.)


SMR

HTS

LTS & K/O Drum

PSA

Furnace

CN 4120: Design II Team 32

Production of Hydrogen via Syngas Route i

TABLE OF CONTENTS

Chapter 1 : PROBLEM DESCRIPTION ........................................................................ 1-1 1.1 PROBLEM STATEMENT FOR PLANT DESIGN SPECIFICATIONS ............ 1-1

1.2 BACKGROUND FOR DEVELOPMENT IN HYDROGEN PRODUCTION .... 1-1

1.2.1 Energy Woes – Away from Fossil Fuels Era ................................................... 1-1

1.2.2 Identifying & Justifying the Production Route – SMR .................................. 1-2

1.2.3 Choice and Significance of Reforming Feedstock – Natural Gas................... 1-4

1.2.4 Steam Methane Reforming (SMR) Reactor .................................................... 1-5

1.2.5 Furnace ............................................................................................................ 1-6

1.2.6 Shift Reactions ................................................................................................. 1-7

1.2.7 Product Purifications ....................................................................................... 1-7

1.2.8 Heat Integration ............................................................................................... 1-8

1.2.9 Cooling Requirements ..................................................................................... 1-9

1.2.10 Use of HYSYS Simulation ........................................................................... 1-10

1.3 REFERENCES ..................................................................................................... 1-10

Chapter 2 STEAM METHANE REFORMER ............................................................... 2-1 2.1 PROBLEM STATEMENT .................................................................................... 2-1

2.1.1 Problem and Specifications ............................................................................. 2-1

2.1.2 Justifications for using SMR ........................................................................... 2-1

2.2 DESIGN METHODOLOGY & PROCESS DESIGN ........................................... 2-2

2.2.1 Outline of Design Methodology ....................................................................... 2-2

2.2.2 Reaction Chemistries ....................................................................................... 2-2

2.2.2.1 Effects of Temperature and Pressure – Revisiting Le Chatelier’s

Principle ............................................................................................................ 2-3

2.2.2.2 Coke Formation, Steam:Methane ratio & Inclusion of CO2 in feed ... 2-3

2.2.3 Choice of Reactor – Tubular Reformer .......................................................... 2-4

2.2.4 Justifications for Choice of Firing Configuration – Side-fired reformer

furnace ...................................................................................................................... 2-5

2.2.5 Justifications for Choice of Fluid Package – Peng-Robinson ......................... 2-6

2.2.6 Choice of Catalyst ............................................................................................ 2-6

2.2.7 Kinetics, Ni-based Catalyst & Role of Support .............................................. 2-7

2.2.8 Justification for Choice of Reactor Inlet Conditions ...................................... 2-7

2.3 PRELIMINARY DESIGN ..................................................................................... 2-8

2.3.1 Establishment of Base Case ............................................................................. 2-8


Production of Hydrogen via Syngas Route ii

2.3.2 Preliminary Simulation using HYSYS ............................................................ 2-8

2.4 DETAILED DESIGN ........................................................................................... 2-10

2.4.1 Development of Critical Profiles via MATLAB & Optimisation ................. 2-10

2.4.2 Design Equations & Key Assumptions.......................................................... 2-10

2.4.3 Fitting into HYSYS Simulation Environment using Plug-Flow Reactor (PFR)

................................................................................................................................. 2-12

2.4.4 Results and Discussions ................................................................................. 2-12

2.4.4.1 Conversion profiles for CH4 and CO2 ............................................... 2-12

2.4.4.2 Temperature and Pressure Variations ............................................... 2-13

2.4.4.3 Component Mole Fractions ................................................................. 2-14

2.4.5 Optimization .................................................................................................. 2-15

2.4.6 Operating Conditions & Streams Conditions ............................................... 2-17

2.5 MATERIALS OF CONSTRUCTION & SIZING .............................................. 2-18

2.5.1 Selection Methodology ................................................................................... 2-18

2.5.2 Justifications for selecting from different grades of stainless steels ............. 2-18

2.5.3 Tube life estimation, Minimum Stress Rupture & Identification of Choice

Material................................................................................................................... 2-19

2.5.4 Sizing – Computation for Tube Thickness .................................................... 2-21

2.5.5 Sizing – Summary .......................................................................................... 2-21

2.6 ECONOMICS & SAFETY CONSIDERATIONS .............................................. 2-22

2.6.1 Economic Analysis (Brief) ............................................................................. 2-22

2.6.2 Safety Consideration for Reactor Design ...................................................... 2-22

2.7 LEARNING & CONCLUSIONS ......................................................................... 2-23

2.8 NOTATIONS ........................................................................................................ 2-24

2.9 FIGURES AND TABLES .................................................................................... 2-25

2.10 ACKNOWLEDGEMENTS................................................................................ 2-25

2.11 REFERENCES ................................................................................................... 2-26

2.12 APPENDIX ......................................................................................................... 2-28

2.12.1 MATLAB Code ............................................................................................ 2-28

2.12.1.1 Main m-file to resolve O.D.E.s .......................................................... 2-28

2.12.1.2 Function m-file to define reactions conditions and O.D.E.s ............. 2-29

2.12.2 List of Equations .......................................................................................... 2-33

2.12.2.1 Rate Equations for reactions and species for 4 O.D.E.s ................... 2-33

2.12.2.2 Mole Fractions for species ................................................................. 2-33


Production of Hydrogen via Syngas Route iii

2.12.2.3 Effectiveness Factors for reactions and species ................................ 2-33

2.12.2.4 Rate & Adsorption constants for reactions 1, 2 and 3 ................................. 2-34

2.12.2.5 Adsorption constants for species .................................................................. 2-34

2.12.2.6 Heat Capacities ............................................................................................. 2-35

2.12.3 Sample Calculations ..................................................................................... 2-35

2.12.4 Typical Natural Gas Compositions ............................................................. 2-35

Chapter 3 : FURNACE .................................................................................................... 3-1 3.1 INTRODUCTION .................................................................................................. 3-1

3.1.1 Furnace design methodology ........................................................................... 3-1

3.1.2 Heat transfer process in fired heater .............................................................. 3-2

3.2 RADIATION ZONE DESIGN ............................................................................... 3-2

3.2.1 Thermal Efficiency of Fired Heater ................................................................ 3-2

3.2.2 Calculation for the number of reformer tubes ............................................... 3-6

3.2.3 Calculation for mass velocity in reformer tubes ............................................. 3-7

3.2.4 Calculation of reformer tube thickness ........................................................... 3-8

3.2.5 Selection of material for reactor tube in radiation section ............................. 3-9

3.2.6 Reformer inner tube diameter....................................................................... 3-11

3.2.7 Furnace layout and design ............................................................................. 3-11

3.2.7.1 Side Fired Heater ................................................................................. 3-11

3.2.7.2 Distance between burners ................................................................... 3-12

3.2.7.3 Burners used at Side Walls ................................................................. 3-13

3.2.7.4 Determination of number of burners .................................................. 3-14

3.2.8 Computations for flue gas temperature ........................................................ 3-15

3.2.8.1 Cold plane area .................................................................................... 3-15

3.2.8.2 Refractory area .................................................................................... 3-15

3.2.8.3 Absorptivity, α ..................................................................................... 3-15

3.2.8.4 Sum of product of area and the absorptivities in the radiant zone .... 3-15

3.2.8.5 Mean beam length ............................................................................... 3-16

3.2.8.6 Partial pressure of CO2 and H2O ....................................................... 3-16

3.2.8.7 Product of partial pressure and mean beam length ........................... 3-16

3.2.8.8 Mean refractory tube wall temperature ............................................. 3-16

3.2.8.9 Two main equations that will be used for iteration to find Tg (flue gas

temp) ................................................................................................................ 3-16

3.2.8.9.1 Radiant zone heat transfer ........................................................... 3-16


Production of Hydrogen via Syngas Route iv

3.2.8.9.2 Radiant zone heat balance ............................................................ 3-16

3.2.8.9.3 Enthalpy of the flue gas as a function of Tg (flue gas temp) ........ 3-17

3.2.8.9.4 Emissitivity of the gas Ф ............................................................... 3-17

3.2.8.9.5 Exchange factor F ......................................................................... 3-17

3.2.9 Residence Time .............................................................................................. 3-18

3.3 CONVECTION SECTION .................................................................................. 3-19

3.3.1 Convection design – Finned tubes ................................................................. 3-19

3.3.2 Design parameters for convection tubes ....................................................... 3-21

3.3.3 Pressure drop in the tubes present in furnace .............................................. 3-22

3.4 STACK DESIGN .................................................................................................. 3-24

3.4.1 Stack diameter ............................................................................................... 3-24

3.4.2 Pressure Drop across stack ............................................................................ 3-24

3.4.2.1 Stack exit loss ....................................................................................... 3-24

3.4.2.2 Frictional Loss in stacks and ducts ..................................................... 3-24

3.4.2.3 Stack entrance loss .............................................................................. 3-25

3.4.2.4 Flue gas pressure drop through the convection section ..................... 3-25

3.4.2.5 Pressure drop at the top of the radiant section .................................. 3-25

3.4.2.6 Pressure gain at the convection section ............................................... 3-25

3.4.3 Stack Height ................................................................................................... 3-26

3.5 MATERIALS FOR CONSTRUCTION OF FURNACE BODY & ADDITIONAL

AUXILIARIES ........................................................................................................... 3-27

3.5.1 Refractory walls ............................................................................................. 3-27

3.5.2 Stack Walls ..................................................................................................... 3-28

3.5.3 Additional auxiliaries ..................................................................................... 3-28

3.5.3.1 Air Preheaters ...................................................................................... 3-28

3.5.3.2 Forced Draft Fan ................................................................................. 3-29

3.5.3.3 Induced Draft Fan ............................................................................... 3-29

3.6 COST ANALYSIS ................................................................................................ 3-30

3.6.1 Purchased Equipment Costs .......................................................................... 3-30

3.6.1.1 Costing for Furnace ............................................................................. 3-30

3.6.1.2 Costing for Air Preheater .................................................................... 3-31

3.6.1.3 Costing for Induced Draft Fan and Forced Draft Fan for Air

Preheating System ........................................................................................... 3-31

3.6.1.4 Burners ................................................................................................ 3-32


Production of Hydrogen via Syngas Route v

3.6.2 Utility Cost ..................................................................................................... 3-32

3.6.2.1 Electricity cost ..................................................................................... 3-32

3.6.3 Total Annual cost ........................................................................................... 3-32

3.7 SUMMARY & CONCLUSION ........................................................................... 3-33

3.8 SPECIFICATION OF FIRED-HEATER ........................................................... 3-34

3.9 REFERENCES ..................................................................................................... 3-35

Chapter 4 : HIGH TEMPERATURE SHIFT REACTOR ............................................. 4-1 4.1 INTRODUCTION .................................................................................................. 4-1

4.1.1 Water gas shift ................................................................................................. 4-1

4.1.2 High temperature shift .................................................................................... 4-2

4.2 PROBLEM DESCRIPTION .................................................................................. 4-3

4.3 REACTION THERMODYNAMICS .................................................................... 4-6

4.3.1 Criteria for Chemical Reaction Equilibrium .................................................. 4-6

4.3.2 Effects of Pressure on Reaction Equilibrium .................................................. 4-7

4.3.3 Effects of Temperature on Reaction Equilibrium .......................................... 4-8

4.4 REACTION KINETICS ...................................................................................... 4-11

4.5 CATALYST .......................................................................................................... 4-12

4.6 REACTOR............................................................................................................ 4-13

4.6.1 Type of reactor ............................................................................................... 4-13

4.6.2 Reactor design ................................................................................................ 4-13

4.7 METHODOLOGY AND CALCULATIONS ...................................................... 4-16

4.7.1 Weight of catalyst .......................................................................................... 4-16

4.7.2 Pressure drop ................................................................................................. 4-19

4.7.3 Thickness of vessel ......................................................................................... 4-24

4.7.4 Reactor size and cost ...................................................................................... 4-24

4.8 HEAT EXCHANGER .......................................................................................... 4-26

4.8.1 Heat Exchanger Design Considerations ........................................................ 4-27

4.8.1.1 Physical properties extraction ........................................................................ 4-27

4.8.1.2 Determination of overall heat transfer coefficient ......................................... 4-28

4.8.1.3 Exchanger type and dimensions ..................................................................... 4-28

4.8.1.4 Heat transfer area ........................................................................................... 4-29

4.8.1.5 Layout and tube size ....................................................................................... 4-29

4.8.1.6 Number of tubes ............................................................................................. 4-29

4.8.1.7 Bundle and shell diameter .............................................................................. 4-29


Production of Hydrogen via Syngas Route vi

4.8.1.8 Tube-side heat transfer coefficient ................................................................. 4-30

4.8.1.9 Shell-side heat transfer coefficient ................................................................. 4-30

4.8.1.10 Overall coefficient ......................................................................................... 4-31

4.8.1.11 Pressure drop ................................................................................................ 4-31

4.9 CONCLUSION ..................................................................................................... 4-32

4.10 NOTATIONS ...................................................................................................... 4-33

4.11 REFERENCES ................................................................................................... 4-35

4.12 APPENDICES .................................................................................................... 4-36

Appendix 4.12.1 .......................................................................................................... 4-36

Appendix 4.12.2 .......................................................................................................... 4-38

Appendix 4.12.3 .......................................................................................................... 4-39

Chapter 5 : LOW TEMPERATURE SHIFT REACTOR .............................................. 5-1 5.1 INTRODUCTION .................................................................................................. 5-1

5.2 LTS DESIGN CONSIDERATIONS ...................................................................... 5-2

5.2.1 Current Status ................................................................................................. 5-2

5.2.2 Kinetics of Low-Temperature Water-Gas-Shift (LTWGS) ........................ 5-3

5.2.2.1 Assumption made for equation (5-3) : .................................................. 5-4

5.2.3 LTS Catalyst .................................................................................................... 5-5

5.2.3.1 Characteristics of the industrial LTS catalyst ...................................... 5-6

5.2.3.2 Preparation ............................................................................................ 5-6

5.2.3.3 Supply .................................................................................................... 5-6

5.2.3.4 Deactivation of LTS Catalyst ................................................................ 5-7

5.2.3.5 LTS catalyst in operation ...................................................................... 5-8

5.2.3.6 Assumptions made for LTS Catalyst .................................................... 5-9

5.2.3.7 Mass balance on the Copper-Zinc catalyst pellet ................................. 5-9

5.2.3.8 Heat balance on the Copper-Zinc catalyst pellet ................................ 5-10

5.2.4 Modeling the converter ............................................................................. 5-12

5.2.4.1 Assumptions made for the converter ................................................ 5-12

5.2.4.2 Reactor mass balance .......................................................................... 5-12

5.2.4.3 Reactor mass balance .......................................................................... 5-13

5.3 DESIGN CONDITIONS ...................................................................................... 5-17

5.3.1 Temperature .................................................................................................. 5-17

5.3.2 Pressure .......................................................................................................... 5-17

5.3.3 Steam to CO ratio .......................................................................................... 5-17


Production of Hydrogen via Syngas Route vii

5.3.4 Design Procedure for LTS outlet compositions and Mass of Catalyst used 5-19

5.3.5 LTS outlet compositions and Mass of Catalysts used ................................... 5-20

5.3.6 Design Procedure for Aspect Ratio ............................................................... 5-21

5.3.7 Results for Aspect Ratio ................................................................................ 5-22

5.3.8 Design Procedure for the dimensions of bed and thickness of vessel wall ... 5-23

5.3.9 Results for bed dimensions and wall thickness ............................................. 5-24

5.3.10 Allowances set for design ............................................................................. 5-25

5.3.11 Study of controlling parameters ................................................................ 5-26

5.4 CHOICE OF A REACTOR BED ........................................................................ 5-27

5.4.1 Cost estimation for the LTS converter .......................................................... 5-28

5.5 DESIGN OF THE KNOCK-OUT DRUM ........................................................... 5-30

5.5.1 Working principle of the knock-out drum .................................................. 5-30

5.5.2 Sizing of the knock-out drum ........................................................................ 5-30

5.5.3 Results and cost estimation ............................................................................ 5-31

LITERATURE REVIEW .......................................................................................... 5-33

CONCLUSION .......................................................................................................... 5-35

BIBLIOGRAPHY ...................................................................................................... 5-36

APPENDIX A1 ........................................................................................................... 5-37

APPENDIX A2 ........................................................................................................... 5-39

APPENDIX A3 ........................................................................................................... 5-39

Chapter 6 : PRESSURE SWING ABSORPTION .......................................................... 6-1 6.1 INTRODUCTION .................................................................................................. 6-1

6.2 PROBLEM STATEMENT .................................................................................... 6-2

6.3 THEORETICAL BACKGROUND ....................................................................... 6-2

6.3.1 Separation via adsorption ................................................................................ 6-2

6.3.2 Pressure-Swing Adsorption (PSA) .................................................................. 6-3

6.3.3 Skarstrom Cycle............................................................................................... 6-3

6.3.4 Adsorbents ....................................................................................................... 6-4

6.4 DESIGN CONSIDERATIONS .............................................................................. 6-5

6.5 ACTUAL MODELING OF PSA ........................................................................... 6-7

6.5.1 Component Mass Balance ............................................................................... 6-8

6.5.2 Overall Mass Balance ...................................................................................... 6-8

6.5.3 Pressure terms.................................................................................................. 6-8

6.5.4 Adsorption rates .............................................................................................. 6-8


Production of Hydrogen via Syngas Route viii

6.5.5 Overall Mass Balance in Dimensionless Form ................................................ 6-9

6.5.6 Component Mass Balance in Dimensionless Form ......................................... 6-9

6.5.7 Dimensionless Pressure terms ......................................................................... 6-9

6.5.8 Dimensionless Langmuir Adsorption Isotherms ............................................ 6-9

6.5.9 Boundary Conditions ..................................................................................... 6-10

6.6 MODEL OPTIMIZATION .................................................................................. 6-11

6.6.1. Process Methodology .................................................................................... 6-12

6.6.2 Initial approximation of the adsorption time from the breakthrough curve... 6-

13

6.6.3 Determination of Cyclic steady state ............................................................. 6-14

6.6.4 Refinement of the pressurization time .......................................................... 6-15

6.6.5 Possible optimization of feed superficial velocity and diameter of the bed . 6-16

6.7 FINAL RESULTS AND DISCUSSIONS ............................................................ 6-18

6.8 COST ESTIMATIONS ........................................................................................ 6-19

6.9 CONCLUSION ..................................................................................................... 6-22

6.10 NOTATIONS ...................................................................................................... 6-23

6.11 APPENDIX ......................................................................................................... 6-25

6.12 CONSTANTS APPLIED IN COMSOL SIMULATION .................................. 6-29

6.13 REFERENCES ................................................................................................... 6-30

Chapter 7 : HEAT EXCHANGER NETWORK ............................................................ 7-1 EXECUTIVE SUMMARY .......................................................................................... 7-1

ACKNOWLEDGEMENTS ......................................................................................... 7-1

7.1 DESIGN METHODOLOGY OF A HEAT EXCHANGER NETWORK ............ 7-2

7.1.1 Determination & Verification of Stream Data Properties Extracted from

Hysys ......................................................................................................................... 7-2

7.1.1.1 Calculations of Maximum Design Velocities ........................................ 7-2

7.1.1.2 Determination of Flow Area Diameter ................................................. 7-3

7.1.1.3 Calculations of Convective Heat Transfer Coefficients (HTC) ........... 7-4

7.1.1.4 Fouling Factors ...................................................................................... 7-5

7.2 TARGETING ......................................................................................................... 7-6

7.2.1 Cost Considerations ......................................................................................... 7-6

7.2.2 Utility Cost Calculations .................................................................................. 7-8

7.2.3 Heat Exchanger Capital Cost Estimations ...................................................... 7-9

7.2.4 Supertargeting ................................................................................................. 7-9


Production of Hydrogen via Syngas Route ix

7.2.5 Comparison between the usage of HP and LP Steam Generation ............... 7-10

7.2.6 Calculation of Utility Targets ........................................................................ 7-10

7.3 MER NETWORK DESIGN ................................................................................. 7-12

7.3.1 Stream matching above pinch ....................................................................... 7-14

7.3.2 Stream matching below pinch ....................................................................... 7-14

7.3.3 Number of units in MER network................................................................. 7-15

7.3.4 Alternative MER Network Designs for Consideration ................................. 7-15

7.3.4.1 Network 1a ........................................................................................... 7-15

7.3.4.2 Network 1b ........................................................................................... 7-16

7.4 NETWORK EVOLUTION .................................................................................. 7-20

7.4.1 Steps involved in network evolution .............................................................. 7-20

7.4.2 Evolution of 1st loop ....................................................................................... 7-20

7.4.3 Evolution of 2nd loop ...................................................................................... 7-21

7.5 HEAT EXCHANGER DESIGN .......................................................................... 7-26

7.5.1 Stream Data ................................................................................................... 7-26

7.5.2 Material of Construction ............................................................................... 7-27

7.5.3 Shell and Tube-Side Fluid Allocation ........................................................... 7-27

7.5.4 Exchanger Type ............................................................................................. 7-28

7.5.5 Baffles ............................................................................................................. 7-28

7.5.6 Tube Dimensions ............................................................................................ 7-28

7.5.7 Tube Arrangements ....................................................................................... 7-29

7.5.8 Calculations .................................................................................................... 7-29

7.5.8.1 Tube-Side Heat Transfer Coefficient Calculations ............................ 7-31

7.5.8.2 Shell-Side Heat Transfer Coefficient Calculations ............................. 7-31

7.5.8.3 Overall Heat Transfer Coefficient Calculations ................................. 7-33

7.5.8.4 Tube-Side Pressure Drop Calculations ............................................... 7-33

7.5.8.5 Shell-Side Pressure Drop Calculations ............................................... 7-34

7.5.9 Modification of Design ................................................................................... 7-34

7.5.10 Exchanger Cost ............................................................................................ 7-35

7.6 RECENT DEVELOPMENTS ................................................................................. 7-36 7.7 HEAT EXCHANGER SPECIFICATION SHEET............................................. 7-37

7.8 INTEGRATED HEN WITH PFD OF PROPOSED HYDROGEN PLANT ..... 7-38

APPENDIX A – STREAM DATA ............................................................................. 7-39

REFERENCE ............................................................................................................. 7-40


Production of Hydrogen via Syngas Route x

Chapter 8 : COOLING TOWER .................................................................................... 8-1 8.1 PROBLEM STATEMENT .................................................................................... 8-1

8.2 WORKING PRINCIPLES OF COOLING TOWER ........................................... 8-2

8.3 Preliminary Design ................................................................................................. 8-3

8.3.1 Selection of cooling tower ................................................................................ 8-3

8.3.1.1 Justification to reject the use of natural draft tower ............................ 8-3

8.3.1.2 Justification to use induced draft tower ............................................... 8-4

8.3.2 Comparison between counter-flow and cross-flow Pattern ........................... 8-4

8.4 DETAILED DESIGN OF COOLING TOWER ................................................... 8-5

8.4.1 Specification of cooling tower design parameters ........................................... 8-5

8.4.1.1 Wet bulb temperature ........................................................................... 8-5

8.4.1.2 Range ..................................................................................................... 8-5

8.4.1.3 Cooling water requirement ................................................................... 8-6

8.4.1.4 Approach ............................................................................................... 8-6

8.4.2 Exit air temperature and water to air flow ratio (L/G) .................................. 8-6

8.4.2.1 Exit air temperature .............................................................................. 8-6

8.4.2.2 Water to air flow (L/G) ratio ................................................................. 8-7

8.4.3 Cooling tower characteristic ............................................................................ 8-7

8.4.4 Loading factor .................................................................................................. 8-8

8.4.5 Dimensions of Tower ......................................................................................... 8-10

8.4.5.1 Fill Height ............................................................................................ 8-10

8.4.5.2 Base area .............................................................................................. 8-10

8.4.5.3 Fill volume ........................................................................................... 8-10

8.4.6 Make-up Water Requirement ....................................................................... 8-11

8.4.6.1 Evaporation loss (E) ............................................................................ 8-11

8.4.6.2 Drift loss (D) ......................................................................................... 8-11

8.4.6.3 Blow-down (B) ..................................................................................... 8-11

8.4.6.4 Makeup water requirement (M) ......................................................... 8-12

8.4.7 Power Requirement ....................................................................................... 8-12

8.4.7.1 Pump power (Pp) .................................................................................. 8-12

8.4.7.2 Fan Power (PF) ..................................................................................... 8-13

8.5 COOLING TOWER INTERNALS ..................................................................... 8-14

8.5.1 Liquid Distributor.......................................................................................... 8-14

8.5.2 Fill ................................................................................................................... 8-15


Production of Hydrogen via Syngas Route xi

8.5.3 Drift Eliminators ............................................................................................ 8-15

8.5.4 Supports ......................................................................................................... 8-16

8.5.5 Cooling tower basin ....................................................................................... 8-16

8.6 MATERIAL OF CONSTRUCTION ................................................................... 8-17

8.6.1 Liquid Distributor.......................................................................................... 8-17

8.6.2 Fills ................................................................................................................. 8-18

8.6.3 Drift eliminator .............................................................................................. 8-18

8.6.4 Mechanical support ....................................................................................... 8-18

8.7 COST ANALYSIS ................................................................................................ 8-19

8.7.1 Construction cost of for cooling tower .......................................................... 8-19

8.7.2 Operating Cost ............................................................................................... 8-20

8.7.2.1 Cost of makeup water .......................................................................... 8-20

8.7.2.2 Cost of Electricity ................................................................................ 8-20

8.7.2 Optimization between the operating and construction cost ......................... 8-21

8.8 ADDITIONAL CONSIDERATIONS TO COOLING TOWER DESIGN ........ 8-22

8.8.1 Water Treatment ........................................................................................... 8-22

8.8.1.1 Corrosion control................................................................................. 8-22

8.8.1.2 Scale control ......................................................................................... 8-23

8.8.1.3 Biological control ................................................................................. 8-23

8.8.2 Environmental Concerns ............................................................................... 8-24

8.9 CONCLUSION ..................................................................................................... 8-25

REFERENCES ........................................................................................................... 8-27

APPENDIX A IMPURITIES FOUND IN COOLING WATER .............................. 8-28

Chapter 9 : ECONOMICS & PROFITABILITY ........................................................... 9-1 9.1 INTRODUCTION .................................................................................................. 9-1

9.2 ASSUMPTIONS ..................................................................................................... 9-1

9.3 CAPITAL COSTS .................................................................................................. 9-2

9.3.1 Computations for Fixed Capital ......................................................................... 9-2

9.3.2 Computations for Total Module Costs ............................................................ 9-7

9.3.3 Computations for Grassroots Costs (FCI) ...................................................... 9-7

9.3.4 Computations for Working Capital ................................................................ 9-8

9.4 MANUFACTURING COSTS ................................................................................ 9-8

9.4.1 Operating labour costs, COL .......................................................................... 9-10

9.4.2 Utility costs, CUT ............................................................................................. 9-11


Production of Hydrogen via Syngas Route xii

9.4.2.1 Electricity ............................................................................................. 9-11

9.4.2.2 Cooling water cost ............................................................................... 9-12

9.4.2.3 Waste treatment costs, CWT ................................................................. 9-12

9.4.3 Raw materials costs, CRM .............................................................................. 9-13

9.4.4 Land lease, CL ................................................................................................ 9-13

9.4.5 Computation of manufacturing costs ............................................................ 9-14

9.4.6 Salvage value .................................................................................................. 9-15

9.4.7 Depreciation ................................................................................................... 9-15

9.4.8 Revenues ......................................................................................................... 9-15

9.5 PROFITABILITY ANALYSIS ............................................................................ 9-16

9.5.1 Land Cost ....................................................................................................... 9-17

9.5.2 After Tax Cash Flow ...................................................................................... 9-17

9.5.2 Rate of Return on Investment (ROROI) ....................................................... 9-18

9.5.3 Net Present Value (NPV) ............................................................................... 9-18

9.5.4 Discounted Cash Flows in Project ................................................................. 9-18

9.6 FEASIBILITY OF STORAGE FACILITIES FOR NATURAL GAS FEED ... 9-21

9.6.1 Capital Costs .................................................................................................. 9-24

9.6.2. Operating Costs ............................................................................................ 9-25

9.6.3 Overall Costs .................................................................................................. 9-25

9.6.4 Economic Compensation ............................................................................... 9-26

9.7 RECOMMENDATIONS ..................................................................................... 9-27

9.8 CONCLUSION ..................................................................................................... 9-29

REFERENCES ........................................................................................................... 9-30

Chapter 10 : SAFETY, HEALTH & ENVIRONMENT (S.H.E.) ................................ 10-1 10.1 INTRODUCTION .............................................................................................. 10-1

10.2 HAZARDS AND OPERABILITY STUDIES (HAZOP) REVIEW ................. 10-2

10.3 PLANT LAYOUT ............................................................................................ 10-23

10.3.1 Segregation ................................................................................................. 10-23

10.3.2 Transportation Considerations ................................................................. 10-24

10.3.3 Administration ........................................................................................... 10-24

10.3.4 Laboratory ................................................................................................. 10-25

10.3.5 Workshop ................................................................................................... 10-25

10.3.6 Control Room ............................................................................................. 10-25

10.3.7 Transformer Substation ............................................................................ 10-26


Production of Hydrogen via Syngas Route xiii

10.3.8 Emergency Services ................................................................................... 10-26

10.3.9 Amenities (Medical Centre and Canteen) ................................................. 10-26

10.3.10 Process and Auxiliary Units .................................................................... 10-27

10.3.10.1 Furnace (Housing SMR) ................................................................ 10-27

10.3.10.2 Reactors (HTS, LTS), PSA and Knockout Drum ......................... 10-27

10.3.10.3 Cooling Tower ................................................................................ 10-28

10.3.10.4 Heat Exchangers ............................................................................ 10-29

10.3.10.5 Flares .............................................................................................. 10-29

10.3.10.6 Wastewater Treatment Plant ........................................................ 10-30

10.4 OCCUPATIONAL SAFETY ........................................................................... 10-31

10.4.1 Personal Protection Equipment (PPE) ...................................................... 10-31

10.4.2 Noise ............................................................................................................... 10-32

10.4.3 Ventilation .................................................................................................. 10-33

10.5 OCCUPATIONAL HEALTH HAZARD IDENTIFICATION ...................... 10-34

10.6 ENVIRONMENTAL IMPACT ASSESSMENT ............................................. 10-38

10.6.1 Objectives ................................................................................................... 10-38

10.6.2 Risk Assessment Matrix............................................................................. 10-38

10.6.3 Elements of Environmental Impact Assessment ....................................... 10-46

10.6.3.1 Gaseous emissions ............................................................................ 10-46

10.6.3.2 Effluent discharge ............................................................................ 10-46

10.6.3.3 Waste management & minimization ............................................... 10-47

10.6.3.4 Energy efficiency ............................................................................. 10-47

10.6.4 Hydrogen Product Life Cycle Assessment ................................................ 10-48

10.6.4.1 Ramifications of Hydrogen LCA .................................................... 10-49

10.7 CONCLUSION ................................................................................................. 10-50

REFERENCES ......................................................................................................... 10-51

Chapter 11 : INSTRUMETNATION & CONTROL.................................................... 11-1 11.1 INTRODUCTION .............................................................................................. 11-1

11.2 PROCESS CONSIDERATION AND DESCRIPTION .................................... 11-2

11.3 PROCESS CONTROL METHODOLOGY ...................................................... 11-3

11.4 SELECTION OF CONTROLLED, MANIPULATED AND MEASURED

VARIABLE ................................................................................................................ 11-4

11.5 DETAILED CONTROL DESIGN FOR REFORMER FEED ......................... 11-5

11.5.1 Steam-to-Methane Ratio Control ................................................................ 11-5


Production of Hydrogen via Syngas Route xiv

11.5.2 Pressure Control Loop for Expander .......................................................... 11-7

11.5.3 Temperature Control Loop to Preheat SMR Feed ..................................... 11-7

11.5.4 Composition Analyzer for SMR Effluent .................................................... 11-8

11.6 DETAILED CONTROL DESIGN FOR SMR FURNACE .............................. 11-9

11.6.1 Air-to-Fuel Ratio Control ............................................................................ 11-9

11.6.2 Temperature Control Loop to Regulate Effluent Exit Temperature ....... 11-10

11.6.3 Pressure Control Loop to Regulate Furnace Draft .................................. 11-11

11.6.4 Flue Gas Exit Temperature Control ......................................................... 11-12

11.6.6 Analyzers for Furnace Control.................................................................. 11-14

11.7 Safety Devices ................................................................................................... 11-14

11.7.1 Pressure Relief Valves................................................................................ 11-14

11.7.2 Process Alarms ........................................................................................... 11-15

11.7.3 Safety Interlocks or Emergency Shutdown System (SIS or ESD) ............ 11-16

11.7.3.1 Implementation of SIS or ESD for the protection of nickel catalyst . 11-

17

11.8 Additional Considerations in Process Control ................................................ 11-18

11.8.1 Redundancy of Air Blowers and Expanders ............................................. 11-18

11.8.2 Isolation Valves and Bypass ...................................................................... 11-19

11.9 REFERENCES ................................................................................................. 11-19

CN 4120: Design II Team 32: Lim Yueh Yang (U046787U) SMR Unit Design

Production of Hydrogen via Syngas Route 1-1

Chapter 1 : PROBLEM DESCRIPTION

1.1 PROBLEM STATEMENT FOR PLANT DESIGN SPECIFICATIONS

This project requires the production of hydrogen in Singapore. This has to be

accomplished via a syngas route, which involves the governing Steam Methane Reforming

(SMR) reactions, as well as the low and high temperature shift reactions. The following

design specifications have been given and the following plant design endeavours to meet

these criteria:

Location of Plant: Singapore

Operation Time: 8000 hours / year

Plant Capacity (PC): 1.25 × 109 m3(STP) / year

Feed Composition (FC) to SMR reactor: 3 : 1 (H2O : CH4)

% CO in H2 Specification at the exit of the shift converter: 0.7%

Purity of hydrogen product: > 99.9% (mole)

Natural Gas Feed: C1 = 97.7%; C2+ = 1.2%; CO2 = 0.7%; N2 = 0.4%

1.2 BACKGROUND FOR DEVELOPMENT IN HYDROGEN PRODUCTION

1.2.1 Energy Woes – Away from Fossil Fuels Era

Recent years saw the rapid developments on alternative energies, in place of their

conventional fossil fuels counterpart. The latter has several disadvantages [R4] associated with

it, including:

(i) Air pollution (formation of NOx, CO & Unburned hydrocarbons contributing to

urban ozone);

(ii) Environmental pollutions (e.g. oil spill during transport)

(iii) Global warming (emission of greenhouse gases) during combustion

(iv) Dependence of fuel supply on oil-producing nations, which could result in

dominance in oil prices



Given these drawbacks, the search for an alternative fuel becomes more pertinent [R4].

One of the possible solutions is the production of hydrogen. In contrast, the latter promises:

(i) Exclusion pollution due to fossil fuels (by-product is H2O & hazards associated

with spills are minimal)

(ii) Exclusion of greenhouse gases

(iii) Removal of price dominance, from the oil-producing nations

(iv) Well-distributed production due to the ease of manufacture.

1.2.2 Identifying & Justifying the Production Route – SMR

The uses of hydrogen extend way beyond the supply for fuels. For instance, hydrogen

could be used in the petrol-chemical industries to make plastics products or it could be used

to produce ammonia in the Haber process. In addition, it has been employed in the refineries

to remove unwanted sulfur contents in crudes via the hydro-de-sulfurization (HDS) units.

Nonetheless, hydrogen does not exist on Earth naturally. To harness of the above-mentioned

uses, a plant has to be designed to produce hydrogen efficiently and safely.

Typically, several methods (Gross, 2005) [R1] are available for hydrogen production.

In the refineries, H2 can be produced in its in-house hydrogen plant or from the CRU

(Catalytic Reformer Unit). H2 produced via the coal gasification route is not aimed at H2

production, rather, it is a by-product of coke production, such as the steel industry in Asia &

Europe [R1]. With more advanced gasification processes, it could also increase the amount of

H2 from coal by a considerable extent. Meanwhile, electrolysis of water promises H2 product

of high purity, but this is dependent on the local costs for electricity. To make it economically

more viable, electricity has to be available at a lower cost. Another instance of using

electricity is the production of Cl2 and NaOH, namely the Chloroalkali process, whereby H2

is produced as a by-product. In fact, more recently, experimental works have gone underway

to produce H2 via photo-electrolysis and biomass gasification.

In this work, one of the most commonly used industrial processes has been adopted,

which is the Steam Methane Reforming (SMR), which accounts for about 45% of world H2

production. This has been illustrated by the following diagram.



Fig 1-1: Distribution of hydrogen production methods in industries

Looking from the perspective of production on an industrial scale, the use of SMR

would provide the economy-of-scale by providing more opportunities for heat

integration (i.e. within SMR unit itself). This reflects a higher degree of optimization for

the usage of utilities. Typically, this can be achieved via steam generation. The latter can be

used for (i) for sale; (ii) for recycle as feed into SMR.

Meanwhile, comparing to other methods (e.g. partial oxidation, auto-thermal

reforming) of syngas production, the SMR route offered the following [R2] competitive

advantages:

Lowest Tprocess required (better cost-savings)

Extensive industrial experience

Best ratio of H2 : CO for production applications of hydrogen

Does not require O2 (cost-savings & safety enhanced)

With such encouraging advantages, the steam reforming process remains as the

most mature and established form of technology to produce hydrogen [R3].

And indeed, several companies world-wide like Haldor-Topsøe, Howe-Baker,

Foster Wheeler, Tokyo Gas Company, McDermott Technology Inc. and IDATech are

employing SMR technology to manufacture hydrogen [R3].



1.2.3 Choice and Significance of Reforming Feedstock – Natural Gas

One of the major factors contributing to the operating characteristics of the reforming

applications is the choice of processing feed for the reformer. In this work, natural gas

(containing predominantly Methane, CH4) has been designated. Upon further research [R3],

few possible reasons for using natural gas include:

Most economic & mature reforming technology

Lower environmental impact (few emissions, except CO2)

Supply of natural gas more readily available

Lower risk of coking (carbon formation)

The following table adapted from literature shows the some of the noteworthy

features for the various choices of reforming feedstock:

Table 1-1: Comparison on different steam reforming feedstock



1.2.4 Steam Methane Reforming (SMR) Reactor

In this design, the SMR unit has been employed to manufacture the product of interest

– Hydrogen (H2). The earlier text has described and justified the need to use SMR and to

have Methane (CH4) as the feedstock. In order to design an efficient reactor to meet the high

CH4 conversion, it is imperative to consider the key factors that played an instrumental role in

influencing the performance of the reactor.

(I) Tube Geometry (related to tube length & diameter, average heat flux & space velocity)

• ↑ Tube Length more economical than ↑ No. of tubes

↑ No. of tubes complicate design at reactor’s inlet and outlet.

• Limit for Tube Length

Threat of tube bending.

Risk of too drastic pressure drop over the catalyst bed.

• ↑ Tube diameter to be accompanied with ↑ Tube Wall thickness

For thinner tubes, ↓ temperature required & better heat transfer (↑cost savings).

Also, less tubes need to be used to meet required conversion.

(II) Firing Configuration (Bottom vs Top vs Terrace vs Side)

• Side-Fired Configuration (with short flames distributed along reactor wall)

Higher level of regulatory control over Tube Wall Temperature.

↑ Design and operational flexibilities.

↑ Average Heat Flux for higher conversion.

Endure more severe reaction conditions.

Lower NOx levels produced in flue gas stream.

• Construction of tubes

Higher level of regulatory control over Tube Wall Temperature.

Creeping strength is a strong function of the choice for tube material

(III) Catalyst (intrinsic activity, surface area, microstructure, porosity, mechanical

resistance, thermal & chemical stability, resistance to carbon deposition)

• Catalyst Structure

Provision of support to give stable micropore system, overcoming sintering

issue when process temperature is above Tamman temperature (Ni: 863K)

Low surface area carriers employed due to high temperatures involved.

Crux: Maximize catalyst activity & heat transfer; Minimize Pressure Drop

Table 1-2: Key design considerations for SMR reactor unit



The SMR reactor typically consists of multiple catalyst-loaded tubes housed within a

furnace. The latter provides the much-needed heat duty due to the characteristic endothermic

reforming reactions, in which CH4 reacts with steam to give the desired H2 product. Given

that the SMR unit represents the heart of the operations for the plant, it is essential for us to

fulfill the key design considerations, as given in Table 1-2.

1.2.5 Furnace

The furnace provides heat to support the endothermic SMR reactions. In this design,

the fuel feed used for the combustion is harnessed from the purge stream of the Pressure

Swing Adsorption (PSA) unit, considering its high H2 (as compared to CH4) content.

However, this purge stream has high carbon dioxide (CO2) content, which does not support

combustion. Consequently, an amine scrubber is also proposed to remove this undesired CO2.

Nonetheless, in the event of insufficient fuel supply by the PSA purge stream, it is

recommended to make up with a natural gas fuel feed. This could originate from the

feedstock of SMR reactor. It is noteworthy that combustion typically occurs at atmospheric

pressure, hence, an expander is to be installed to decrease the pressure of the SMR natural gas

feedstock, before allowing the fuel to proceed to the burner.

To demonstrate the advantages conferred by the side-fired configuration, small

premix burners would line up along the walls of the straight wall furnace, as such burners

provide short flame length and ease for temperature control. The flue gas generated from the

combustion process carries a net amount of heat for which is transported upward to the

convection section through the use of induced draft fan. The heat carried by the flue gas is

then used to heat up the process streams passing through the convection section.

In the convection section, steam is generated within the tubes closest to the radiation

section. This is followed by two other process streams, namely, SMR feed and combustible

air. Steam generation is situated closest to the radiation section because heat transfer is most

efficient for heat exchange between two different phases. Through effective process control

and instrumentation, the process variables within the furnace are kept constant. This helps to

maintain the product yield, while keeping the operating environment safe.



1.2.6 Shift Reactions

The main purpose of the high temperature shift reactor is to derive more H2 from the

one of the SMR products, carbon monoxide (CO). The feed into the High Temperature Shift

(HTS) reactor is at a relatively high level of CO as it exits from the steam methane reformer.

This CO was reacted with steam to form more H2 with the following water gas shift equation:

222 HCOOHCO +⇔+ molkJH rxn /447.44−=∆

There is a need to couple the high temperature shift reaction with a low temperature

shift (LTS) reaction because of the exothermic nature of the water gas shift reaction.

Therefore, high conversion occurs at low temperatures. However, the rate of reaction is too

slow (i.e. compromised) at low temperatures. Thus, the HTS reactor is employed to ensure a

high reaction rate, while its LTS counterpart maintains the required conversion.

Based on iron oxide as the catalyst, the design of the HTS was able to convert a

13.34% CO feed, to 3.0% CO, after which it is fed into the LTS. The designed conversion of

the high temperature shift reactor was 75.27%. Due to the adiabatic reaction in high

temperature shift, the temperature of the feed was raised from 627K to 692K.

The feed was cooled to 493K prior to entry into the low temperature shift reactor,

which employed the Copper-Zinc Oxide catalyst supported on alumina. A CO conversion

efficiency of 82.9% was obtained, which corresponds to a 0.5mol% CO (dry basis) in the

outlet stream of the LTS. This product stream was then transferred to the knock-out drum,

where liquid water was separated from the other gaseous products. The latter then flowed to

the PSA columns for further purification.

1.2.7 Product Purifications

For this design, pressure swing adsorption (PSA) was adopted as the preferred mode

of purification due to the stated requirement of attaining 99.9% in product purity, which

otherwise was not achievable through the conventional use of a CO2 scrubber and a

methanator (95–97%).



The high concentration of hydrogen (75% at the entrance) could lead to possible

hydrogen embrittlement, thus the material of construction chosen was carbon steel clad with

stainless steel as this material possessed a lower material factor as compared to stainless steel,

which translated to a lower bare module cost (1.8 vs 3).

Through the prior installation silica gel, the amount of water entering PSA after

exiting the knockout drum was assumed to be negligible. Thus the chosen adsorbents for PSA

were activated carbon and zeolite 5A. Activated carbon was utilized to remove hydrocarbons,

such as CH4, C2+

and CO2 due to the preferential adsorption isotherms that these components

exhibit with activated carbon. Similarly, zeolite 5A was employed to remove CO and N2.

According to the Polybed design, this comprises of 7-10 beds with the incorporation of

various operation steps, such as pressurization, high pressure adsorption, blowdown and

purge, a final product of 99.9% purity and 85% hydrogen recovery was attainable.

Subsequently, the PSA tail gas was routed to the furnace as a source of fuel for combustion.

1.2.8 Heat Integration

Energy integration involves the usage of process streams within the plant itself to

fulfil the heating and cooling requirements at various points of the process. An optimal

solution would be of utmost importance in a chemical plant, as this would help to mitigate the

rising cost of utilities associated with increased fuel cost. Therefore, to achieve optimal heat

integration, the systematic development of a heat exchange network (HEN) would have to be

carried out.

The usage of a HEN would be an integral step in the maximization of energy recovery.

The use of pinch analysis would be critical in lessening the requirements for hot and cold

utilities, which are major components of the operating cost of the hydrogen plant. However,

this must be balanced with the increased capital investment associated with the installation of

heat exchangers.

It was found that a furnace was required to provide the necessary heat of reaction for

SMR in normal operation. Preliminary calculations showed that the large amount of heat

produced by the furnace provided for the entire heating requirement in the hydrogen plant.



This results in the lack of need for hot utilities, i.e. a threshold problem ensued with

only cold utilities required. Hence, possible cold utilities to be considered would be the

generation of high or low pressure steam, and the usage of cooling water to cool low-grade

heat. This could result in the presence of a utility pinch, which would be tackled by a similar

application of pinch analysis, treating the utility stream as a dummy process stream. A

multitude of network variations and possible evolution of the network would also be

considered to obtain the most economical and practical solution for the energy integration of

the designed hydrogen plant.

1.2.9 Cooling Requirements

The main purpose of the cooling tower is to reject the low grade heat absorbed from

process stream into the atmosphere by means of latent heat of evaporation and sensible heat

transfer. The cooling tower in this hydrogen plant is designed to provide a continuous flow of

cooling water required for the condensation and elimination of water vapor in the outlet

stream of the LTS reactor, before it is fed into the PSA columns for purification of H2 and

removal of CO2.

The design of the cooling tower is based on an induced draft counter-flow

configuration. This is because this type of configuration does not experience any recirculation

which can cause a drop in cooling tower efficiency due to higher wet bulb temperature and in

the long run, it is more economical due to lower power requirement for auxiliary units such

as fans and pumps.

In the design of this cooling tower, it is assumed by heuristic that the maximum inlet

temperature of cooling water to be 120˚F and cooling water exit temperature to be 90˚F and

the ambient wet bulb temperature is derived from the average daily maximum dry bulb

temperature and mean humidity. Hence, the performance of the cooling tower can be

optimized by manipulating the exit air temperature and it is found to be 105˚F, which is the

average of the inlet and outlet water temperature.



1.2.10 Use of HYSYS Simulation

The simulation of the hydrogen plant was performed in Hysys. Care must be

exercised in the selection of the fluid package of choice as any unsuitability would be

reflected in the obtainment of inaccurate simulation results.

Thus Peng-Robinson (PR) Equation Of State (EOS) was adopted as the preferred fluid

package. AspenTech recommended it for oil, gas and petrochemical applications due to its

special enhancement in HYSYS for the generation of accurate phase calculations over a wide

range of operating conditions (T > -271°C, P < 1000kPa). Our reaction conditions were well

within the range. Furthermore, literature values obtained for the reactor units had been based

primarily on the PR EOS.

The PSA was reflected as a component splitter in the PFD. PSA was a process unit

that could not be adequately simulated in Hysys, thus its simulation was performed in

COMSOL. The target specifications for the various major units had been met with the

convergence of the Hysys simulations, which also implied an overall satisfactory plant design.

1.3 REFERENCES

[R1]: Tom Gross. (2005). Hydrogen – An Overview. Foundation for Nuclear Studies Briefing.

[R2]: Wilhelm, D., Simbeck, D., Karp, A., Dickenson, R. (2001). Syngas production for gas-

to-liquids applications: technologies, issues and outlook. Fuel Proc. Tech., Vol 71 – P139

[R3]: Ferreira-Aparicio, P., Benito, M. J. & Sanz, J. L. (2005). New Trends in Reforming

Technologies: from Hydrogen Industrial Plants to Multifuel Microreformers. Catalysis

Reviews, 47:4, P491-588.

[R4]: Marshall Brain. How the Hydrogen Economy Works. Adapted on 15th Apr 2008 from:

http://auto.howstuffworks.com/hydrogen-economy.htm



Chapter 2 STEAM METHANE REFORMER

2.1 PROBLEM STATEMENT

2.1.1 Problem and Specifications

In this report, a Steam Methane Reformer (SMR) reactor unit is to be designed. The

plant is to produce hydrogen via the syngas route. The SMR reactor is one of the first units in

the process stream, and hence its design would be critical for the downstream process units,

in a bid to achieve an overall economical, safe and efficient plant for the hydrogen

production.

Amongst all, the design specifications for Team 32 are shown as follow:

• Location of Plant: Singapore

• Operation Time: 8000 hours / year

• Plant Capacity (PC): 1.25 × 109 m3(STP) / year

• Feed Composition (FC) to SMR reactor: 3 : 1 (H2O : CH4)

• % CO in H2 Specification at the exit of the shift converter: 0.7%

• Purity of hydrogen product: > 99.9% (mole)

• Natural Gas Feed: C1 = 97.7%; C2+ = 1.2%; CO2 = 0.7%; N2 = 0.4%

2.1.2 Justifications for using SMR

Justifications to leverage upon the SMR reactor unit for hydrogen production have been

found in literature. For instance, Wilhelm et. al. (2001) [R8] described the following

advantages, which are aligned with the current intention of the usage of the SMR unit. These

advantages have made SMR the chosen reforming concept. Hence, this project endeavours to

produce hydrogen via this syngas route.

Lowest Tprocess required (better cost-savings)

Extensive industrial experience

Best ratio of H2 : CO for production applications of hydrogen

Does not require O2 (cost-savings & safety enhanced)

CN 4120: Design II Team 32: Lim Yueh Yang (U046787U) SMR Unit Design Report


2.2 DESIGN METHODOLOGY & PROCESS DESIGN

2.2.1 Outline of Design Methodology

Fig 2.2.1a Flowchart to illustrate design methodology

2.2.2 Reaction Chemistries

CH4 + H2O ↔ CO + 3 H2 Eqn (2-1)

CO + H2O ↔ CO2 + H2 Eqn (2-2)

CH4 + 2 H2O ↔ CO2 + 4 H2 Eqn (2-3)

3 governing equations responsible for the reactions in the reactor are given as above.

At this point, it is crucial to note that Beurden (2004) [R24] described that Eqn (2-3) is not a

combination of the Eqn (2-1) and Eqn (2-2) as CO2 is produced in both Eqn (2-2) and (2-3),

implying that the latter itself does not represent an overall reaction.



2.2.2.1 Effects of Temperature and Pressure – Revisiting Le Chatelier’s Principle

La Chatlier’s Principles can be used to explain the effects of the operating conditions

for a typical SMR. A higher inlet temperature (typically 723 – 923K [R25]) would drive the

endothermic reactions (Eqn (2-1) & (2-3)) forward to produce more H2 product.

This is in contrast with that of the water-gas- shift (WGS) reaction (Eqn (2-2)), which

is favoured at lower temperature and not affected by pressure (same molar ratio on both sides

of reaction no volume expansion).

Meanwhile, the stoichiometries of these 2 reforming reactions also suggested that

forward reactions are favoured when a lower pressure is used. This is to allow for volume

expansion to occur since the number of moles of product is greater than that of reactants.

2.2.2.2 Coke Formation, Steam:Methane ratio & Inclusion of CO2 in feed

Also, the Steam:Methane ratio (sc) used is 3. This coincides with what is typically

found in industrial practices, which suffice in suppressing coke formation [R24] during the

reaction. The presence of the carbon deposits during coke formation is detrimental to the

process as it would result in tube blockage, forming hot spots that can very well destroy the

tubes, threatening both the economics and safety of the process. Since this SMR reactor unit

design does not consider formation of coke, the choice of sc = 3 is made during the start of

the project to favour the design considerations of not involving coking as one of the reactions.

The suppression of coke formation is further promoted by the inclusion of CO2 [R25]

in the feed gas (Boudouard reaction during coking: 2CO = C + CO2), as mentioned in the

design brief. This shifts the Boudouard reaction backwards and thereby suppressing coke

formation. In addition, adding CO2 at the inlet of the reformer helps to save on hydrocarbon

feedstock and decrease the H2:CO formed in the SMR product stream. With these advantages

in mind, in industrial practice, some of these CO2 are typically being recycled from the SMR

product stream or being imported from another source.



2.2.3 Choice of Reactor – Tubular Reformer

Nielsen (1993) indicated that the steam-reforming reaction typically involves catalysts

being loaded into tubes, which are in turn housed in a furnace to satisfy the highly endo-

thermic reaction. The tube material has to be capable to withstand the high temperature and

the temperature gradient (e.g. 1223 K at outlet [R25]).

As such, these tubular reactors typically experienced very huge stresses. Given that

upper limit of the tolerable stress value for the tubes is very much affected by the maximum

tube wall temperature and heat flux, a small rise in the maximum tube wall temperature could

very well resulting a reduction of life expectancy for these tubes.

Typical average lifespan of these reformer tubes can be around 100,000 hours. Given

that the current plant is designed to run at 8000 hours/yr, this would allow use for up to

a good 12.5 years. Such tubular reformer would be choice reactor for the current design

because it allows catalysts

Fig 2.2.3a: Typical Natural Gas & Reformer Catalysts. Retrieved from Midrex on World

Wide Web: http://www.midrex.com/uploads/documents/Catalyst(1)1.pdf



Fig 2.2.3b: Tubular reformer configurations [R23, R25]

2.2.4 Justifications for Choice of Firing Configuration – Side-fired reformer furnace

The side-fired heating configuration is chosen because this has:

Provided greater degree of control for Ttube wall to allow a more robust operation, to meet

the demands of the production by enduring more severe operating conditions. Also, a

higher average heat flux of 313800 kJ/h/m2 [R26] can be allowed.

Shorter residence time discouraged [R25] formation of nitrogen oxides (NOx), up to

<50ppm, which is critical for the current design since it is assumed that no NOx is formed

in the SMR unit, even though there is presence of some N2 in the feed gas and that the

reactor is to be operating at high temperatures.



2.2.5 Justifications for Choice of Fluid Package – Peng-Robinson

The Peng-Robinson fluid package is chosen because it gives a better prediction of

liquid densities than the Soave-Redlich-Kwong equation of state [R19]. In fact, the Peng-

Robinson equation of state is the most commonly used for systems containing non-polar

components [R20]. From the Table 2.2.4a below, it can be determined that Peng Robinson fluid

package remains the best choice as the system [R21].

Table 2.2.4a: Recommended

Property Package based on type

of system

2.2.6 Choice of Catalyst

Catalyst for steam reforming is typically nickel-based, which is cheaper than its noble

metal counterparts. Key criteria (besides costing) are to maximise activity and heat

transfer, while not causing a huge pressure drop. Non-metallic options are yet to be made

commercial because of the inherent low catalyst activity and thereby the impact of pyrolysis

[R27]. The current work is based on that of Rajesh et. al. (2000) but there is no clear mention

of the catalyst used. Hence, the catalyst would follow that used by Xu and Froment (1989)

[R3,4], which is mentioned by Rajesh et. al, namely Ni/MgAl2O4. In this case, note that the

stability of the MgAl2O4 catalyst support is also very crucial in order to withstand conditions

during operation, start-up and shut-down [R27].



2.2.7 Kinetics, Ni-based Catalyst & Role of Support

To facilitate the reforming reactions, the Ni-based catalyst is added to lower the

activation energies requirements. Intrinsic activity of the catalyst depends [R25] on the catalyst

surface available for chemisorption to occur. To do so, the cleavage of a C–H is needed, and

this entails overcoming the barrier of 52 kJ/mol and contacting a free Ni adsorbent site with

free neighbouring site. This chemisorption is of CH4 plays an instrumental role in

determining the reaction rates of the SMR process (typically given by CH4=CH3* CH2*

CH*=C*). Indeed, this is also aligned [R25] with the typical 1st order kinetics mentioned by

many [R2,3,4,5].

Nonetheless, as observed from Eqn (2-17) in Chapter 2.12.2.1, the presence of the

equilibrium adsorption constant for water (KH2O) may explain for a negative reaction order

with respect to the feed steam, giving rise to a ‘retarding effect’ whose extent may vary with

the choice of catalyst. In view of that, it is thus important to include alkali or use of magnesia

as support for the Ni catalyst, which provide for an enhanced level of adsorption of steam

molecules to avoid carbon formation.

2.2.8 Justification for Choice of Reactor Inlet Conditions

Rajesh et. al. (2000) [R1] suggested that inlet temperature for the SMR reactor should

not be lower than 725 K due to thermodynamic limitations, which thus prevent possible

formation of gum on the reformer catalyst. This would block the catalyst surface [R24] via

polymerisation of the adsorbed CnHm radicals. Such progressive deactivation is possible since

species like ethane (C2H6) is one of the components in the natural gas feed for the current

project. Meanwhile, operating at temperatures higher than 900 K is also not probable

because of the limitations for the maximum amount of heat energy that can be harnessed

form the flue gas. Meanwhile, it is also suggested [R1] that the inlet pressure is typically

between 2400 – 3000 kPa, after accounting for the normal pressures at which H2 is to be

produced, and the presence of natural gas as the feed. With these in mind, the arithmetic

averages for the inlet temperature and pressure (812.5 K and 2700 kPa) are used during the

detailed design of the SMR reactor unit.



2.3 PRELIMINARY DESIGN

2.3.1 Establishment of Base Case

The governing reactions, key reactions and products for the reactor are first identified.

From the specified plant capacity provided in the design brief, and given the conversion

specified in literature (see trailing paragraph), a preliminary mass balance is then performed

across the reactor unit, after obtaining the SMR outlet flow rates from mass balance for the

inlet of trailing High Temperature Shift (HTS) unit. This gave an initial value for the flow

rates for the feed to be used.

Research efforts have been invested to find the conversion of the major component,

namely Methane (CH4). Coincidentally, for a Steam:Methane ratio of 3, Moulijn et. al. (2001)

[R9] has reported on the expected methane slip is to be 21% (assuming P = 2700 kPa (27

bar); T = 1123.15 K), corresponding to a 79% (i.e. 100% - 21%) CH4 conversion.

Fig 2.3.1a: Literature data Moulijn et. al.

(2001) [R9] to support conversion obtained

during preliminary design is valid at the

assumed conditions (2700kPa & 1123.15K)

2.3.2 Preliminary Simulation using HYSYS

Initial efforts for SMR reactor design has been made via data from Hou & Hughes

(2001) [R6] literature data and using the Equilibrium Reactor module in HYSYS as a

simulation tool.



Accordingly, the reactors were stainless steel tubes with 0.01m internal diameter and

0.338m long. Catalyst used was ICI steam reforming catalyst 57-4, which is of a cylindrical

fashion with 4 axial holes and provided by ICI Katalco. These catalysts have been crushed

prior to their usage. Amongst the various reaction equations given, the 3 predominant ones

haven seen selected and their respective equilibrium constants have been given.

Reaction Reaction Equilibrium Rate constant, Ki

1) CH4 + H2O = CO + 3H2 1.198 × 1017 exp(-26830/T) Eqn (2-4)

2) CO + H2O = CO2 + H2 1.767 × 10-2 exp(4400/T) Eqn (2-5)

3) CH4 + 2H2O = CO2 + 4H2 2.117 × 1015 exp(-22430/T) Eqn (2-6)

Table 2.3.2a: Reaction Equilibrium Rate constants, from Hou & Hughes (2001) [R6]

For a preliminary estimate, assume if TSMR, Exit = TSMR, Equilibrium = 1123.15K, then:

K1 = 5.058 × 106 (kPa)2; K2 = 8.884 × 10-1; K3 = 4.484 × 106 (kPa)2

With these Ki expressions, an Equilibrium Reactor module is specified in the

HYSYS Simulation Environment. Inlet pressure is assumed to be 270kPa (arithmetic average

of the TSMR, Inlet from the Rajesh et. al. [R1]), while inlet temperature is set to be 923.15K [R??].

Further research depicts an allowable pressure drop of 200 kPa [R27] may be typically used

across the SMR reactor unit. Overall, the simulation has achieved a CH4 conversion of ≈≈≈≈

79 mol%, which corresponds to that stipulated by the literature, as mentioned in

Chapter 2.3.1. The resultant process flow conditions are then presented in the interim report.

The latter would be used as an initial estimate for the respective downstream units.

Note: Streams and reactors data obtained at this stage from the HYSYS Process Flow

Diagram (PFD) via use of the Equilibrium Reactor) module merely provide initial iteration

values for solving of MATLAB Ordinary Differential Equations (O.D.E.s) for the detailed

design.



2.4 DETAILED DESIGN

2.4.1 Development of Critical Profiles via MATLAB & Optimisation

To materialise the design of the reactor in detail, MATLAB is employed to solve for

the O.D.E.s given in literature [R1], after obtaining initial iteration values from the preliminary

design. The intention is to develop critical profiles and useful data via literature data and

assumptions, and fit the findings into the HYSYS Simulation Environment (using a PFR

module) for the overall plant design.

The MATLAB code developed could be found in the Appendix (Chapter 2.12.1). 2

m-files are written. The first one (Chapter 2.12.1.1) specified the reaction inlet conditions and

the tube dimensions, while using ode15s to resolve the O.D.E.s along the axial direction. The

second m-file (Chapter 2.12.1.2) specified the 4 O.D.E.s to be solved, as well as how each

parameter is obtained from literature.

Optimization is then performed by manipulating the various parameter like heat flux,

number of tubes and tube inner diameter being used. Since bulk of the written code followed

closely to the O.D.E.s developed by Rajesh et. al. (2000) [R1], the arithmetic averages for the

inlet pressure (2700 kPa) and temperature (812.5 K) of the range recommended by the same

journal are used to be the base case in this case.

2.4.2 Design Equations & Key Assumptions

Rajesh et. al. [R1] summarised the kinetic model and energy balance of the steam

reforming process, using a side-fired configuration with Mg-supported Ni catalyst. The 4

major O.D.E.s to be resolved via MATLAB are shown as follow.

0;)4

(0

2

4

444 ===t

CH

CHCHiCH

F

rRd

dt

dχ

ηπχ CH4 Differential Mass Balance Eqn (2-7)

0;)4

(0

2

2

222 ===tCO

COCOiCO

F

rRd

dt

dχ

ηπχ CO2 Differential Mass Balance Eqn (2-8)

itgps

s PPcatbedvoidD

catbedvoidG

dt

dP=

−−=

=03;

)(

)1(75.1

ρφ Momentum Balance Eqn (2-9)



iti

iiib

i

F

overallmass

TTrHd

H

CpGdt

dT=

∆−+==

=∑ 0

3

1

;))(41

ηρ Energy Balance Eqn (2-10)

There 4 O.D.E.s have to be resolved to observe the profiles of how the CH4

conversion, CO2 conversion, pressure and temperature vary along the length of the reactor (t).

Expanded form of the expressions for the rates, effectiveness factors, adsorptions and rates

constants are available in Chapter 2.12.2.

Several key assumptions are made, including:

1) Constant heat flux <HF> is assumed, which implied [R28] that a well-controlled firing rate

is assumed. This is more probable [R27] in side-fired furnace used in this work due to

greater degree of adjustments and control over the tube wall temperature. Consensus is

also reached by Furnace Team to use a constant heat flux for the current work.

2) Effectiveness factor < iη > for the ith reactions, which vary along the tube length, are fitted

with Excel, whose coefficients of the polynomial equations are then input into MATLAB.

3) There is negligible (0.001H2:1CH4) H2 recycle <hc> from the Pressure Swing Adsorption

(PSA) unit to the reformer feed.

4) Feed ratios <sc, dc & nc> and the total mass flow rate are obtained from HYSYS PFD.

The Adjust Function is used for more accurate values of the component molar flow rates.

5) No formation of NOx due to use of short flames of the side-fired furnace (Chapter 2.2.4).

6) No coking due to presence of CO2 in the reformer feed (Chapter 2.2.2).

7) N2 and higher carbons (C2+ = C2H6 in this work) remained inert in the SMR unit. Typical

compositions of Natural Gas feed may be found in Chapter 2.12.4 of the Appendix.

8) Catalyst dimensions and characteristics follow that of Rajesh et. al. (2000) [R1].

9) Density of process gas < gρ > varies along the tube length <t>, which is taken to be the

division of the mass flow rate over volumetric flow rate.

10) Volumetric flow rate is taken to essure

eTemperaturtconsgasUniversalrateflowMolar

Pr

tan ××

for any axial position <t>, hence Ideal Gas Law is assumed here.

11) For the MATLAB code being written, expressions for C2 is not available in the literature

used [R1], hence this is lumped the other inert species N2. However, in the HYSYS

Simulation, these two species are distinguished.



2.4.3 Fitting into HYSYS Simulation Environment using Plug-Flow Reactor (PFR)

The HYSYS PFR module is employed to simulate the findings from MATLAB by:

1) Obtaining conditions of inlet and outlet streams, and also tubing and catalyst

specifications [R1,2,3,4] from optimisation in MATLAB, which are then fitted into the

HYSYS Simulation Environment.

2) Upon convergence in HYSYS, initial flow rates are then fitted back into MATLAB to re-

generate the critical profiles. These profiles are to be aligned with that in HYSYS.

As the flow fashion is now being modelled as plug flow [R7] in HYSYS, it is assumed

that no axial mixing occurs. This coincides with the intention of generating 1-D critical

profiles (with respect to Length of Reactor, t). Also, a Heterogeneous Catalytic reaction set

is chosen since the SMR reactor unit involves In the current HYSYS simulation, the PFR is

being segmented to 50, instead of the default value of 20. Hesketh (2003) [R7] described

that the increased number of steps conferred higher accuracy when resolving the O.D.E.s,

since now more steps used to resolve the O.D.E.s.

2.4.4 Results and Discussions

2.4.4.1 Conversion profiles for CH4 and CO2

Fig 2.4.4.1a:

Conversion

Profiles from

MATLAB



A desirable conversion of up to 80% is achieved for methane (CH4). This is in close

agreement with results obtained in the preliminary findings (Chapter 2.3.2), as well as that in

literature (Chapter 2.3.1). This also affirms that the SMR unit design is able to meet its

conversion targets for H2 production, as proposed in the interim report.

2.4.4.2 Temperature and Pressure Variations

Fig 2.4.4.2a: Pressure and Temperature profiles along length of reformer tube

Nielsen (1993) [R25] reported that typical outlet can be as high as 1223 K (9500C).

Hence, the current SMR exit temperature of 1100 K is still lower than the literature value.

As mentioned earlier in Chapter 2.2.3, it is noteworthy that neither very high temperature

nor great temperature gradient is encouraged since this may increase stress on the

reformer tubes, which greatly reduces the lifespan of the tubes.

Meanwhile, the pressure drop is about 65 kPa, which is lower than the proposed drop

of 200 kPa (2 bar) in the interim report. A lower pressure drop would mean that

downstream compressions could be avoided/minimised [R26], resulting in cost savings.



2.4.4.3 Component Mole Fractions

Fig 2.4.4.3a: Component Mole Fractions from MATLAB

Fig 2.4.4.3b: Component Mole Fractions from HYSYS



Comparing Fig 2.4.4.3a & b with

that of Fig 2.4.4.3c, it can be

observed that the simulated profiles

from both MATLAB and HYSYS

PFR are aligned with that in

literature by Rajesh et. al. (2001)

[R12] on multi-objective optimization.

This may signified that the current

chosen design configuration has also

been optimized.

Fig 2.4.4.3c: Graph from literature

on mole fraction profiles [R12]

2.4.5 Optimization

Optimization has been performed to obtain the desirable reformer tube configuration

and operating conditions. At this juncture, since costs of the various materials (e.g. catalyst,

tube materials) are typically proprietary information, hence, the chosen tube dimensions and

the conditions of operation is based on other parameters. Through MATLAB, graphs of

increasing and

decreasing a

particular

parameter (e.g.

inner diameter, di)

are painstakingly

plotted. One

example is as

shown in Fig

2.4.5a.

Fig 2.4.5a: Effect of varying tubular inner diameter on CH4 conversion profile



Fig 2.4.5b: Effect of varying pressure on CH4 conversion profile

Table 2.4.5a: Effect of Manipulating Parameters on Critical Profiles (from MATLAB)

As illustrated above, trends observed by increasing and decreasing selected

parameters are summarised in Table 2.4.5a. After several rounds of optimizing and fine

adjustments, the operating conditions and stream properties are presented in Chapter 2.4.6.



2.4.6 Operating Conditions & Streams Conditions

Table 2.4.6a: Stream Properties & Operating Conditions (HYSYS PFR & MATLAB)

For overall integration purposes and in view that both HYSYS & MATLAB values

are in close agreement, these values from HYSYS are passed down to the downstream units.

Note that from HYSYS, a heating duty of 5.949 ×××× 108 kJ/hr is required.



2.5 MATERIALS OF CONSTRUCTION & SIZING

2.5.1 Selection Methodology

Selection of reactor materials has been made via the following considerations [R15].

• Conditions of exposure for reactor equipment being determined.

• Availability of materials being explored.

• Suitable material being identified.

• Ensured choice of material being substantiated with certifications.

The reactions involved exposed the SMR reactor unit to high temperature and

pressure. Comparisons between operating conditions for using MATLAB and HYSYS PFR

module revealed that these are typically as high as 1130 K (15750F) and 2700 kPa (391 psi)

[R16]. In addition, there are chances of hot gas corrosion due to the high mass velocity.

2.5.2 Justifications for selecting from different grades of stainless steels

In view of these conditions, stainless steel is a suitable material for the construction of

these SMR reactor tubes, which can be summarised as such:

Higher C Content Offers greater creep resistance than other metals

Addition of Ni and Cr Resistance to carburization and creep being enhanced.

Different grades of heat resistance steel, namely, HH, HK, HD and HF. HK have been

specifically found to be of great use for SMR due to their high creep and rupture strength

even up to 1150°C. Most importantly, they offer resistance to hot gas corrosion. Literature

[R16] revealed typical material of choice for SMR is HK40. Fig 2.5.2a illustrates the relative

tensile strengths of the different stainless steel grades.



Fig 2.5.2a: Relative tensile strengths of different stainless steel grades

2.5.3 Tube life estimation, Minimum Stress Rupture & Identification of Choice Material

From Fig 2.5.3a [R16], at the high inlet pressure of 2700kPa (391 psi) and at the high

outlet temperature of 1130 K (1575 °F) for the reactor, the furnace tube life is found to

more than 20.3 years. This provides a confirmation that the material can be employed as

suitable for use in steam methane reforming processes.

Meanwhile, Fig 2.5.3b [R16] illustrates the minimum stress to rupture for the HK40

material as compared to other grades. The figure implies a lower performance of HK40 grade

compared to HP grades. However, since the maximum temperature is 1130 K, which is low

compared to the maximum temperature at which these stainless steel grades can withstand,

the pressure factor is taken for higher consideration in selection of the suitable metal type.

Since HK40 is capable of withstanding high pressure, HK40 would thus be chosen as the

choice of material for the reformer tubes (assuming economic considerations are ignored).



Fig 2.5.3a: Estimation of furnace tube life for a given set of operating temperature & pressure

Fig 2.5.3b: Minimum Stress to rupture for chosen material



2.5.4 Sizing – Computation for Tube Thickness

HK40 (25% Cr, 20% Ni) are usually used for steam reformer tubes [R16]. It has been

chosen for its high creep and rupture strengths. It is also resistant to hot gas corrosion and

hence is usually employed in steam methane reforming processes [R15]. Meanwhile, it

provides creep resistance up to 980°C (1253 K) [R17], making it a good candidate for the

current design whereby the highest temperature 856.85°C (1130 K). A standard code formula

is employed here to compute the minimum wall thickness required [R18], for sizing purpose

and specification of the PFR module in the HYSYS Simulation Environment.

m

in

min

kPa

psikPa

ksi

psiksi

inm

inm

kPa

psikPa

PES

FCAd

P

ta

i

wall

2

min

1001.1

1

0254.0399.0

1

1448.027006.085.0

1

10002.3

039.00254.0

1

2

127.0

1

1448.02700

6.0

2

−×=

×=

××+××

+××

×=

−

+

=

P = Max Pressure (to be in psig) = 2700 kPa (highest pressure at inlet);

di = Inner Diameter (to be in inches) = 0.127 m;

FCA = 10-year corrosion allowance (to be in inches) = 0.039 in [R18];

Sa = Minimum Creep Stress for HK40 (to be in psi) = 3.2 ksi;

E = Weld Efficiency Factor = 0.85 [R17].

Hence, a tube wall thickness of 1.01××××10-2 m would be used.

2.5.5 Sizing – Summary

Inner Diamter (di) : 0.127 m (5”) Outer Diameter (do) : 0.147 m

Wall thickness : 0.0101 m No. of tubes needed : 450 (HK-40 Steel)

Further details of sizing of SMR reactor are to be done with Furnace unit counterpart

since these reformer tubes are housed in the furnace itself. These would then be presented in

the Final Team Report.



2.6 ECONOMICS & SAFETY CONSIDERATIONS

2.6.1 Economic Analysis (Brief)

Table 2.6.1a: Costing Analysis for SMR unit [R29,30,31,32]

Due to lack of credible literature data to support the costing analysis, an extremely

rough estimate is given above in Table 2.6.1a. Computation efforts are done to illustrate how

costing analysis can be done if there is access to proprietary pricing information while

working as a real engineer. Nonetheless, further research effort would be done and cost

estimations (with other considerations) would be put forth in the Final Team Report.

2.6.2 Safety Consideration for Reactor Design

Reactor is the heart of the plant design. Given the high speed steam and natural gas to

be fed into the reactor, and the huge amount of heat is needed to supply to this endothermic

reforming process, a great deal of safety consideration has to be in place to ensure that the

plant and its operators can operate safety and efficiently.

In the Final Team Report where Process & Instrumentation analysis is done, more

findings on safety considerations would be reported, with collaborative efforts with the

furnace counterpart.



2.7 LEARNING & CONCLUSIONS

In this report, the tubular reformer used for to produce hydrogen via a syngas route is

designed. Both preliminary and detailed design achieved similar conversions of the major

reactant component, namely Methane (CH4). Also, the detailed design is done via resolving

Ordinary Differential Equations in MATLAB by obtaining information from literature

research and making key assumptions with appropriate justifications. Besides, integration

into the team’s overall process flow diagram has also been done by inserting values from the

MATLAB model (which takes into account of both intrinsic kinetics and diffusional

limitations) into a Plug-Flow-Reactor module in the HYSYS Simulation Environment.

Results are considerably satisfactory since both values from MATLAB and HYSYS are in

close agreement with each other, and aligned with that found in literature. Meanwhile, tube

dimensions and material of constructions, brief economic analysis and safety considerations

have also been covered in this report

Several learning can be derived from the current work of design a Steam Methane

Reforming (SMR) reactor unit. Besides the need to plough through several literature data, the

author has learnt to exercise discretion when researching through the available information,

via perform the Principles of Chemical Engineering taught earlier.



2.8 NOTATIONS

Sym. Description Units

Ki Equilibrium rate constant for ith rxn K1 & K3: (kPa)2; K2: unitless Kj Equilibrium adsorption constant for specie ‘j’ KCH4,H2,CO: (kPa)-1;

KH2O: unitless T Temperature Kelvin (K) P Pressure kPa ki Rate constant for ith rxn k1,3: (kmol•kPa0.5)/(kg•h)

k2: (kmol•kPa-1)/(kg•h) yj Mole Fraction for specie ‘j’ molj / moltotal

jχ Conversion for specie ‘j’ j = CH4, CO2

sc Molar feed ratio of H2O:CH4 unitless hc Molar feed ratio of H2(recycle stream):CH4 unitless dc Molar Feed ratio of CO2:CH4 unitless nc Molar Feed ratio of N2:CH4 unitless ri Rate of ith rxn at catalyst surface kmol/(h•kgcat) rj Rxn rate for specie ‘j’ at catalyst surface j = CH4, CO2; kmol/(h•kgcat)

iη Effectiveness factor: ith rxn unitless

jη Effectiveness factor: Conversion for jth specie j = CH4, CO2; unitless

Cpmean,i Mean Specific Heat Capacity of ith rxn kJ/(kmol•K) Cpoverall Overall Specific Heat Capacity of ith rxn kJ/(kmol•K) t Axial position in reformer tube m di Inner diameter of reformer tube m do Outer diameter of reformer tube m catbedvoid Catalyst bed void fraction 0.605 [R1] Gs Mass velocity of process gas (from HYSYS) kg/(h•m2) Gmass Mass velocity of process gas (from HYSYS) kg/(s•m2)

sφ Sphericity of catalyst pellet 0.6563 [R1]

pD Equivalent length for catalyst pellet 0.0174131m [R1]

bρ Bulk density of catalyst 1362.0 kg/m3 [R1]

gρ Density of process gas at any axial position kg/m3 [R1]

F Reformer feed rate kmol/h R Sum of all molar ratios in feed unitless HF Heat Flux (= Heat Transfer Coefficient × Temp.

Diff b/w Tinner tube wall & Touter tube wall) kcal/(h•m2) Assume to be: 25000 Btu/h/ft2 or 283913.167 kJ/h/m2

- iH∆ Heat of the ith reaction kcal/kmol

iRH,

∆ Heat of ith reaction kcal/kmol

iυ Stoichiometric coefficient for ith reaction ‘-’ reactants; ‘+’ products



2.9 FIGURES AND TABLES

• Fig 2.2.1a Flowchart to illustrate design methodology

• Fig 2.2.3a: Typical Natural Gas & Reformer Catalysts

• Fig 2.2.3b: Tubular reformer configurations

• Table 2.2.4a: Recommended Property Package based on type of system

• Fig 2.3.1a: Literature data to support conversion obtained during preliminary design is

valid at the assumed conditions (2700kPa & 1123.15K)

• Table 2.3.1a: Reaction Equilibrium Rate constants, from Hou & Hughes (2001)

• Fig 2.4.4.1a: Conversion Profiles from MATLAB

• Fig 2.4.4.2a: Pressure and Temperature profiles along length of reformer tube

• Fig 2.4.4.3a: Component Mole Fractions from MATLAB Fig 4.4.3b: Component Mole

Fractions from HYSYS

• Fig 2.4.4.3c: Graph from literature on mole fraction profiles

• Fig 2.4.5a: Effect of varying tubular inner diameter on CH4 conversion profile

• Fig 2.4.5b: Effect of varying pressure on CH4 conversion profile

• Table 2.4.5a: Effect of Manipulating Parameters on Critical Profiles (from MATLAB)

• Table 2.4.6a: Stream Properties & Operating Conditions (HYSYS PFR & MATLAB)

• Fig 2.5.2a: Relative tensile strengths of different stainless steel grades

• Fig 2.5.3a: Estimation of furnace tube life for a given set of operating temp. & pressure

• Fig 2.5.3b: Minimum Stress to rupture for chosen material

• Table 2.6.1a: Costing Analysis for SMR unit

2.10 ACKNOWLEDGEMENTS

This section dedicates acknowledgements to all who have helped the author by

offering their valuable insights and advices. In particular, the author would like to express

gratitude to Prof. Kawi for his advice, as well as to Mr Thanneer for his consultation on the

MATLAB codes and functions. Last but not least, this work would not have been possibly

done without the coordination and teamwork in Team 32 (CN4120 – Sem 2 of AY2007/08),

hence the author would like to thank all of them for their assistance and understanding.



2.11 REFERENCES

[R1]: J.K. Rajesh, Santosh K.Gupta, G.P.Rangaiah & Ajay K. Ray. (2000). Multiobjective

Optimization of Steam Reformer Performance Using Genetic Algorithm. Ind. Eng. Chen.

Res.: Vol 39 – P706-717.

[R2]: S.S.E.H. Elnashaie & S.S. Elshishini. (1993). Modelling, Simulation And Optimization

Of Industrial Fixed Bed Catalytic Reactors. Gordon And Breach Science Publishers.

[R3]: Jianguo Xu & Gilbert F. Froment (1989). Methane Steam Reforming, Methanation and

Water-Gas Shift: I. Instrinsic Kinetics. AIChE Journal: Vol 35 – No.1

[R4]: Jianguo Xu & Gilbert F. Froment (1989). Methane Steam Reforming: II. Diffusional

Limitations and Reactor Simulation. AIChE Journal: Vol 35 – No.1

[R5]: Kaihu Hou & Ronald Hughes. (2001). The kinetics of methane steam reforming over a

Ni/α-Al2O catalyst. Chemical Engineering Journal: Vol 82 – P311-328.

[R6]: Kaihu Hou & Ronald Hughes. (2001). The kinetics of methane steam reforming over a

Ni/α-Al2O catalyst. Chemical Engineering Journal.: Vol 82 – P311-328.

[R7]: Robert P. Hesketh. (2003). Catalytic Rates & Pressure Drops in PFR Reactors: HYSYS

3.0.

[R8]: Wilhelm, D., Simbeck, D., Karp, A., Dickenson, R. (2001). Syngas production for gas-

to-liquids applications: technologies, issues and outlook. Fuel Proc. Tech., Vol 71 – P139

[R9]: Moulijn, J., Makkee, M., van Diepen, A. (2001). Chemical Process Technology. John

Wiley & Sons Ltd (England).

[R10]: J.A. Moulijn, A.E. van Diepen & F. Kapteijn. (2001). Catalyst deactivation: is it

predictable? What to do? Applied Catalysis A: General 212 – P3-16.

[R11]: Chang Samuel Hsu & Paul R. Robinson. (2006). Practical Advances in Petroleum

Processing. Springer Science+Business Media, Inc.

[R12]: J. K. Rajesh, S. K. Gupta, G. P. Rangaiah & A. K. Ray. (2001). Multi-objective

optimization of industrial hydrogen plants. Chemical Engineering Science: Vol56–P999-1010.

[R13]: J.M. Smith, H.C. Van Ness & M. M. Abbott. (2005). Introduction to Chemical

Engineering Thermodynamics – 7th Edition. McGraw-Hill International Edition.

[R14]: Dilton, C.P. (1992). Materials selection for the chemical process industries. McGraw-

Hill.

[R15]: Retrieved on 16th March 2008 from World Wide Web:

http://www.valve-world.net/pdf/11022.pdf.



[R16]: V. Ganapathy. (1982). Applied heat transfer. PennWell.


http://www.fitness4service.com/publications/pdf_downloads/Jaske-Shannon%20Paper.PDF


http://www.kubotametal.com/alloys/heat_resistant/HK-40.pdf


http://www.tu-harburg.de/vt2/pe2000/Dokumentation/PE2000_Kap7A1.htm

[R20]: Chorng H. Twu, John E. Coon & David Bluck. (1997). A Comparison of the Peng-

Robinson and Soave-Redlich-Kwong. Equations of State Using a New Zero-Pressure-Based

Mixing Rule for the Prediction of High Pressure and High Temperature Phase Equilibria.

Simulation Sciences Inc.


http://che.sut.ac.ir/People%5CCourses%5C65%5CCHEM_2_3.PDF


http://encyclopedia.airliquide.com/Encyclopedia.asp?GasID=41

[R23]: Kelly Ibsen. (2006). Equipment Design and Cost Estimation for Small Modular

Biomass Systems, Synthesis Gas Cleanup, and Oxygen Separation Equipment. Nexant Inc.

[R24]: P. van Beurden. (2004). On the Catalytic Aspects Of Steam-Methane Reforming.

[R25]: J.R. Rostrup-Nielsen. (1993). Production of synthesis gas. Catalysis Today: Vol 18.

[R26]: Ib Dybkjaer. (1995). Tubular reforming and authothermal reforming of natural gas –

an overview of available processes. Fuel Processing Technology: Vol 42 – P85-101.

[R27]: H.I.deLasa, G.Dogu & A.Ravella. (1991). Chemical Reactor Technology for

Environmentally Safe Reactors and Products. Applied Sciences: NATO ASI Series Vol. 225

[R28]: J.R.Rostrup-Nielsenn, L.J.Christiansen & J.H.Bak Hansen. (1988). Activity of Steam

Reforming Catalysts: Role and Assessment. Applied Catalysis: Vol43–P287-303.

[R29]: Price of Nickel and Magnesium. Retrieved from World Wide Web on 19th March

2008: http://www.sciencelab.com/page/S/PVAR/10-807

[R30]: Price of HK-40 alloy (approximate): Retrieved from World Wide Web on 19th March

2008: http://www.meps.co.uk/Stainless%20Prices.htm

[R31]: Price of Al2O4: Retrieved from World Wide Web on 19th March 2008:

http://www.encyclopedia.com/doc/1G1-104622322.html

[R32]: Density of HK40 alloy: Retrieved from World Wide Web on 19th March 2008:

http://sg.search.yahoo.com/search?p=density+of+HK40+alloy&fr=yfp-t-

web&toggle=1&cop=&ei=UTF-8



2.12 APPENDIX

2.12.1 MATLAB Code

2.12.1.1 Main m-file to resolve O.D.E.s

clear all clc close all format long global catbedvoid phis Dp tube_length rhob rhos Pi Ti num_tubes_total di di_inch do global Ac sc hc dc nc R %SMR Inlet Conditions================================================= Pi = 2700; %kPa; Inlet Pressure Ti = 812.5; %K; Inlet Temperature; r/f [R1], this is b/w 725K and 900K %Tubing Dimensions & Number=========================================== num_tubes_total = 450; %TOTAL number of tubes di_inch = 5; %inch; Specify di in inches di = di_inch*0.0254; %m; Inner tube diameter Ac = pi*(di^2)/4; %m^2; Tube Cross-Sectional Area tube_length = 11.95; %m; Length of tube %Catalyst and Bed properties============================================= Dp = 0.0174131; %m; Pellet equivalent diameter catbedvoid = 0.605; %unitless; Catalyst bed void fraction rhob = 1362.0; %kg/m^3; Catalyst bed density rhos = 2355.2; %kg/m^3; Solid catalyst density phis = 0.6563; %unitless; Pellet sphericity %Molar Feed Compositions & Ratios======================================= sc = 3; % steam/CH4 molar feed ratio ==> FIXED hc = 0.0001; % H2/CH4 molar feed ratio; H2 from PSA RECYCLE dc = 0.00716496; % mol.CO2 / mol.CH4; from HYSYS PFR nc = 0.004094094; % mol.N2 / mol.CH4; from HYSYS PFR %Specify conditions and solve for the 4 ODES================================= tspan = [0 tube_length]; [t,y] = ode15s('smrodes',tspan,[0,0,Pi,Ti]); %Simulation results==================================================== figure subplot(2,2,1) hold on plot(t,y(:,1)) xlabel('Length of Reactor, m'); ylabel('CH4 conversion, xCH4'); subplot(2,2,2) hold on plot(t,y(:,2)) xlabel('Length of Reactor, m'); ylabel('CO2 conversion, xCO2');



subplot(2,2,3) hold on plot(t,y(:,3)) xlabel('Length of Reactor, m'); ylabel('Pressure, kPA'); subplot(2,2,4) hold on plot(t,y(:,4)) xlabel('Length of Reactor, m'); ylabel('Temperature, K');

2.12.1.2 Function m-file to define reactions conditions and O.D.E.s

function dy = smrodes(t,y) % Only declare those used as CONSTANTS to be "global" global R sc hc dc nc num_tubes_total Ac HF HF_btuperft2perhr do di di_inch rhob catbedvoid phis Dp global Fo FN2o FN2 G global MWCH4 MWH2O MWH2 MWCO MWCO2 MWN2 MWC2H6 FN2 FN2o global deltaA1 deltaB1 deltaC1 deltaD1 global deltaA2 deltaB2 deltaC2 deltaD2 global deltaA3 deltaB3 deltaC3 deltaD3 global H10 H20 H30 %These computations are done according to [R1], if there exists deviations, %these would be typically be mentioned as comments. %This m-file computes the 4 ODES to be resolved, whose solutions are then %input in the matrix y, so as to resolve them al xCH4 = y(1); %CH4 molar conversion at any axial position xCO2 = y(2); %CO2 molar conversion at any axial position P = y(3); %Pressure at any axial position T = y(4); %Temperature at any axial position %Computation for R=================================================== R = 1 + sc + hc + dc + nc; %sum of molar feed ratios %Adsorption constants for Individual Species============================== KCH4 = (6.65*10^(-6)).*exp(4604.28./T); %kPa^-1 KH2O = (1.77*10^(3)).*exp(-10666.35./T); %unitless; r/f [R1] & [R5] KH2 = (6.12*10^(-11)).*exp(9971.13./T); %kPa^-1 KCO = (8.23*10^(-7)).*exp(8497.71./T); %kPa^-1 %Equilibrium constants for Rxn I, II & III================================ K1 = 10266.76.*exp(-26830./T + 30.11); %kPa^2 K2 = exp(4400.0./T - 4.063); %unitless K3 = K1.*K2; %kPa^2 %Rate Coefficients for Rxn I, II & III==================================== k1 = 9.490*10^16.*exp(-28879./T); %kmol.kPa^0.5/kg.h k2 = 4.390*10^4.*exp(-8074.3./T); %kmol.kPa^-1/kg.h k3 = 2.290*10^16.*exp(-29336.0./T); %kmol.kPa^0.5/kg.h



%Effectiveness Factors for Rxn I, II & III================================ n1 = (-7*10^-7).*t.^6 + (3*10^-5).*t.^5 - 0.0004.*t.^4 + 0.0029.*t.^3 - 0.0102.*t.^2 + 0.015.*t + 0.0165; %excel ==> deg 6 if (t <= 3.4) n2 = -0.0059.*t.^5 + 0.0559.*t.^4 - 0.1971.*t.^3 + 0.3263.*t.^2 - 0.2316.*t + 0.0889; %excel ==> deg 5 elseif (t == 3.4) n2 = 0; else n2 = (-7*10^-6).*t.^4 + 0.0004.*t.^3 - 0.0074.*t.^2 + 0.0651.*t - 0.2158; %excel ==> deg 6 end n3 = (-6*10^-7).*t.^6 + (3*10^-5).*t.^5 - 0.0004.*t.^4 + 0.0033.*t.^3 - 0.0132.*t.^2 + 0.0229.*t + 0.0102; %excel ==> deg 6 %GASEOUS Mole Fraction Basis for ALL SPECIES============================ yCH4 = (1-xCH4)./(R+2.*xCH4); yH2O = (sc-xCH4-xCO2)./(R + 2.*xCH4); yCO = (xCH4 - xCO2)./(R+2.*xCH4); yCO2 = (dc + xCO2)./(R + 2.*xCH4); yH2 = (hc + 3.*xCH4 + xCO2)./(R+2.*xCH4); yN2 = nc./(R+2.*xCH4); %Molecular Weights for ALL SPECIES===================================== MWCH4 = 16.043; %kg/kmol; from [R2] MWH2O = 18.01524; %kg/kmol; from [R2] MWH2 = 2.016; %kg/kmol; from [R2] MWCO = 28.01; %kg/kmol; from [R2] MWCO2 = 44.01; %kg/kmol; from [R2] MWN2 = 28.0134; %kg/kmol; from [R2] %Flowrates (Mass & Molar); Density of Process Gas========================= Fo = 1.040E4/num_tubes_total; % kmol/h %Reformer Molar Feed Flow Rate at inlet for ONE TUBE; from HYSYS FN2o = 10.4; % kmol(N2)/h %yN2o * Fo = FN2o (for ONE TUBE) where yN2o is initial N2 mole fraction; from HYSYS FN2 = FN2o; % kmol(N2)/h %FN2o is the N2 molar flow rate at inlet = FN2 is the N2 molar flow rate at any axial length F = FN2./(yN2*num_tubes_total); % kmol/h %Total molar flow rate at any axial length for ONE TUBE sv = ((F.*8.314.*T)./P)/(Ac); % m/h %Superficial Velocity = Volumetric Flow Rate (ASSUME Ideal Gas) / Ac G = 183287.978554998/num_tubes_total; % kg/h %Total mass flow rate for ONE TUBE; from HYSYS rhog = G./((F.*8.314.*T)./P); % kg/m^3; %Density of gas mixture = Mass flow rate per tube / Volumetric Flow Rate per tube Gmass = rhog.*sv; % kg/h/m^2; Mass velocity in per HOUR basis, to be used for dy(4) Gs = Gmass/3600; % kg/s/m^2; mass velocity in per SECOND basis, to be used for dy(3)



%Rate of Rxns at catalyst surface at any axial position kmol/h/kgcat================== E = 1 + P.*(KCO.*yCO + KCH4.*yCH4 + KH2.*yH2) + KH2O.*yH2O./yH2; r1 = (k1./(E.^2.*yH2.^2.5.*P.^0.5)).*(yCH4.*yH2O-(P.^2.*yH2.^3.*yCO./K1)); r2 = ((k2.*P)./(E.^2.*yH2)).*((yCO.*yH2O)-((yH2.*yCO2)./K2)); r3 = (k3./(E.^2.*yH2.^3.5.*P.^0.5)).*(yCH4.*yH2O.^2-(((yH2.^4.*yCO2).*(P.^2))./(K1.*K2))); rCH4 = r1+r3; rCO2 = r2+r3; nch4 = (n1.*r1 + n3.*r3)./(r1+r3); %effectiveness factor for adsorption of CH4 nco2 = (n2.*r2 + n3.*r3)./(r2+r3); %effectiveness factor for adsorption of CO2 %Specific Heat Capacities=============================================== %Compute constants to find Cp(mean) for Rxn I, II & III [R13] %Recall Rxn1: CH4 + H2O = CO + 3H2 deltaA1 = (-1)*(1.702)+(-1)*(3.470)+(1)*(3.376)+(3)*(3.249); deltaB1 = (-1)*(9.081/10^3)+(-1)*(1.450/10^3)+(1)*(0.557/10^3)+(3)*(0.422/10^3); deltaC1 = (-1)*(-2.164/10^6) +(-1)*(0)+(1)*(0)+(3)*(0); deltaD1 = (-1)*(0)+(-1)*(0.121/10^-5)+(1)*(-0.031/10^-5)+(3)*(0.083/10^-5); %Recall Rxn2: CO + H2O = CO2 + H2 deltaA2 = (-1)*(3.376)+(-1)*(3.470)+(1)*(5.457)+(1)*(3.249); deltaB2 = (-1)*(0.557/10^3)+(-1)*(1.450/10^3)+(1)*(1.045/10^3)+(1)*(0.422/10^3); deltaC2 = (-1)*(0)+(-1)*(0)+(1)*(0)+(1)*(0); deltaD2 = (-1)*(-0.031/10^-5)+(-1)*(0.121/10^-5) +(1)*(-1.157/10^-5) +(1)*(0.083/10^-5); %Recall Rxn3: CH4 + 2H2O = CO2 + 4H2 deltaA3 = (-1)*(1.702)+(-2)*(3.470)+(1)*(5.457)+(4)*(3.249); deltaB3 = (-1)*(9.081/10^3)+(-2)*(1.450/10^3)+(1)*(1.045/10^3)+(4)*(0.422/10^3); deltaC3 = (-1)*(-2.164/10^6) +(-2)*(0)+(1)*(0)+(4)*(0); deltaD3 = (-1)*(0)+(-2)*(0.121/10^-5) +(1)*(-1.157/10^-5) +(4)*(0.083/10^-5); %Note that 1.987 is multiplied to convert kJ/mol.K to kcal/kmol.K %Also, note that 298.15 K is reference temperature %Also, T/298.15 is tile in [R13] Pg. 141 Eqn (4-20) Cpmean1 = 1.987*(deltaA1 + (deltaB1/2)*(298.15).*(T/298.15+1) + (deltaC1/3).*((298.15)^2).*((T/298.15).^2 + (T/298.15) + 1) + deltaD1./((T/298.15).*298.15^2)); Cpmean2 = 1.987*(deltaA2 + (deltaB2/2)*(298.15).*(T/298.15+1) + (deltaC2/3).*((298.15)^2).*((T/298.15).^2 + (T/298.15) + 1) + deltaD2./((T/298.15).*298.15^2)); Cpmean3 = 1.987*(deltaA3 + (deltaB3/2)*(298.15).*(T/298.15+1) + (deltaC3/3).*((298.15)^2).*((T/298.15).^2 + (T/298.15) + 1) + deltaD3./((T/298.15).*298.15^2)); %Note that divide by 4.196 such that kJ/kmol ==> kcal/kmol H10 = (2.061*10^5) /(4.186); %ENDOTHERMIC; [R1] H20 = (-4.11*10^4) /(4.186); %EXOTHERMIC; [R1] H30 = (1.650*10^5) /(4.186); %ENDOTHERMIC; [R1] %Heats of Reactions=================================================== H1 = H10 + Cpmean1.*(T-298.15); %kcal/kmol H2 = H20 + Cpmean2.*(T-298.15); %kcal/kmol if (n2<0) H2 = -H2; %kcal/kmol else H2 = H2; %kcal/kmol end



H3 = H30 + Cpmean3.*(T-298.15); %kcal/kmol sum1 = -H1.*n1.*r1; %for dy(4) sum2 = -H2.*n2.*r2; %for dy(4) sum3 = -H3.*n3.*r3; %for dy(4) %Heat Flux (CONSTAN VALUE IS ASSUMED HERE; r/f to Chapter 2.4.2 for explanation) HF_btuperft2perhr = 25000; %btu/h/ft^2; from [R11] HF = HF_btuperft2perhr*2.71427502; %kcal/h/m^2; %Weight Fraction for ALL SPECIES======================================= %From ABOVE ==> G is Total mass flow rate for ONE TUBE; from HYSYS wCH4 = (yCH4.*F)*MWCH4./G; wH2O = (yH2O.*F)*MWH2O./G; wH2 = (yH2.*F)*MWH2./G; wCO = (yCO.*F)*MWCO./G; wCO2 = (yCO2.*F)*MWCO2./G; wN2 = (yN2.*F)*MWN2./G; %Specific Heat Capacities for ALL SPECIES================================= CpCH4 = 1.987*(1.702 + (9.081/10^3).*T + (-2.164/10^6).*T.^2 + (0).*T.^-2)/MWCH4; %kcal/kg.K CpH2O = 1.987*(3.470 + (1.45/10^3).*T + (0/10^6).*(T.^2) + (0.121/10^-5).*(T.^-2))/MWH2O; %kcal/kg.K CpH2 = 1.987*(3.249 + (0.422/10^3).*T + (0/10^6).*(T.^2) + (0.083/10^-5).*(T.^-2))/MWH2; %kcal/kg.K CpCO = 1.987*(3.376 + (0.557/10^3).*T + (0/10^6).*(T.^2) + (-0.031/10^-5).*(T.^-2))/MWCO; %kcal/kg.K CpCO2 = 1.987*(5.457 + (1.045/10^3).*T + (0/10^6).*(T.^2) + (-1.157/10^-5).*(T.^-2))/MWCO2; %kcal/kg.K CpN2 = 1.987*(3.280 + (0.593/10^3).*T + (0/10^6).*(T.^2) + (0.040/10^-5).*(T.^-2))/MWN2; %kcal/kg.K Cpoverall = wCH4.*CpCH4 + wH2O.*CpH2O + wH2.*CpH2 + wCO.*CpCO + wCO2.*CpCO2 + wN2.*CpN2; %kcal/kg.K %Ordinary Differential Equations========================================== dy(1) = Ac.*R.*rhob.*nch4.*rCH4./Fo; dy(2) = Ac.*R.*rhob.*nco2.*rCO2./Fo; dy(3) = -(1.75*(Gs^2)*(1-catbedvoid))./(phis*Dp*((catbedvoid)^3).*rhog)/1000; %/1000 is to account to change Pa to kPa so as to use P for other functions dy(4) = (1./(Gmass.*Cpoverall)).*((4*HF)/di + rhob.*(sum1+sum2+sum3)); % note that units for Cpoverall is kcal/kg.K dy = dy'; %Note: This is skeletal MATLAB developed to solve for the 4 O.D.E.s. Additional strings of

code used to generate the various plots shown in Chapter 2.4.4, 2.4.5 and 2.4.6 are not shown

here due to space constraint. In general, just need to comment off the ‘clear all’ command,

and then vary the parameter(s) (e.g. num_tubes_total), and record the figures in a new matrix

after each run. Thus, graphs with different num_tubes_total can then be plotted on one plot.



2.12.2 List of Equations

2.12.2.1 Rate Equations for reactions and species for 4 O.D.E.s

cat

COH

OHCH

Hkgh

kmol

K

yyPyy

PyE

kr

•−= );(

1

32

5.05.22

11

2

24

2

Eqn (12-11)

cat

COH

OHCO

Hkgh

kmol

K

yyyy

yE

kr

•−= );(

22

22

22

2

2

Eqn (12-12)

cat

COH

OHCH

Hkgh

kmol

K

Pyyyy

PyE

kr

•−= );(

3

24

2

5.05.32

33

22

24

2

Eqn (12-13)

314rrrCH += Eqn (2-14) 322

rrrCO += Eqn (2-15)

ncdchcscR ++++= 1 (used in Eqn (4-1) & (4-2)) Eqn (2-16)

2

2

22244)(1

H

OH

OHHHCHCHCOCOy

yKyKyKyKPE ++++= Eqn (2-17)

2.12.2.2 Mole Fractions for species

4

4

4 2

1

CH

CH

CHR

yχ

χ

+

−= Eqn (2-18)

4

24

2 2 CH

COCH

OHR

scy

χ

χχ

+

−−= Eqn (2-19)

4

24

2 CH

COCH

COR

yχ

χχ

+

−= Eqn (2-20)

4

2

2 2 CH

CO

COR

dcy

χ

χ

+

+= Eqn (2-21)

4

24

2 2

3

CH

COCH

HR

hcy

χ

χχ

+

++= Eqn (2-22)

4

2 2 CH

NR

ncy

χ+= Eqn (2-23)

2.12.2.3 Effectiveness Factors for reactions and species

321 &, ηηη are obtained via fitting polynomials using Microsoft Excel. Points are

specified via identifications of coordinates for these 3 curves via vigorous read-off.

0165.0015.0)0102.0()0029.0()0004.0()103()107( 2345567

1 ++−+−×+×−= −− ttttttη

Polynomial of degree 6 Eqn (2-24)

0889.02316.0)3263.0()1971.0()0559.0()0059.0( 2345

2 ++−−+−= tttttη

Polynomial of degree 5 (for t < 3.4) Eqn (2-25)



2158.0.00651.0)0074.0()0004.0()107( 2346

2 −+−+×−= − ttttη

Polynomial of degree 4 (for t =3.4 and t > 3.4) Eqn (2-26)

0102.00229.0)0132.0()0033.0()0004.0()103()106( 2345567

3 ++−+−×+×−= −− ttttttη

Polynomial of degree 6 Eqn (2-27)

31

3311

4 rr

rrCH +

+=

ηηη Eqn (2-28)

32

3322

2 rr

rrCO +

+=

ηηη Eqn (2-29)

2.12.2.4 Rate & Adsorption constants for reactions 1, 2 and 3

hkg

kPakmol

Tk

••

−×=

5.06

1 ;0.28879

exp10490.9 Eqn (2-30)

hkg

kPakmol

Tk

••

−×=

−14

2 ;3.8074

exp10390.4 Eqn (2-31)

hkg

kPakmol

Tk

••

−×=

5.016

3 ;0.29336

exp10290.2 Eqn (2-32)

2

1 ;11.30)0.26830(

exp76.10266 kPaT

K

+−

= Eqn (2-33)

unitlessT

K ;063.4)0.4400(

exp2

−−−

= Eqn (2-34)

2

213 ; kPaKKK ×= Eqn (2-35)

2.12.2.5 Adsorption constants for species

16 ;)28.4604(

exp1065.64

−−

−−×= kPa

TKCH Eqn (2-36)

unitlessT

K OH ;)35.10666(

exp1077.1 3

2

−×= Eqn (2-37)

111 ;)13.9971(

exp1012.62

−−

−−×= kPa

TK H Eqn (2-38)

17 ;)71.8497(

exp1023.8 −−

−−×= kPa

TKCO Eqn (2-39)



2.12.2.6 Heat Capacities

)(,, RimeaniRi TTCpHH −+∆=∆ Eqn (2-40)

( ) ( ) ( )( ) ( )( )

+

++

+

−+=

2

0

0

0

2

0

2

0

0

0, 1

31

2T

T

T

D

T

T

T

TTC

T

TTBARCp iiiiii

iiimean

υυυυ

Values of A, B, C and C for the respective species for the ith reaction are found in [R13].

Eqn (2-40)

2.12.3 Sample Calculations

Most of the calculations are performed via MATLAB and the HYSYS Simulation

Environment, so long as the relevant parameters are specified. Hence, sample calculations

would not be shown in this work. All the MATLAB written have comment statements

intended to make the code self-explanatory.

2.12.4 Typical Natural Gas Compositions

Figure is retrieved from Midrex from World Wide Web on 19th March 2008 at:

http://www.midrex.com/uploads/documents/Catalyst(1)1.pdf.

CN 4120: Design II Team 32: Ng Su Peng (U046929L) Furnace Unit Design Report


Chapter 3 : FURNACE

3.1 INTRODUCTION

3.1.1 Furnace design methodology

Furnaces serve various purposes in process industries such as column reboilers,

reactor-feed preheaters etc. Unlike typical furnaces; the furnace used for the steam methane

reforming has additional design considerations compared to conventional furnaces. The

furnace in this project is used to provide energy for the steam methane reaction. It provides

single-phase/multiple-component heating. In addition, the convection section of the furnace

serves to extract excess heat from the flue gas to heat up process streams from other part of

the plant. Typically, 70% of the heat generated by the burner goes to the radiation section

while the remaining 30% goes to the convection. Single phase multiple components heating

will be carried out by the furnace. Catalysts would be placed in the reactor tubes lining the

refractory.

The furnace design would incorporate the following considerations:

(1) Capacity and size of furnace (2) Dimensions of reactor tubes

(3) Material selection (4) Safety considerations

Fig 3.1.1a: Furnace design methodology



3.1.2 Heat transfer process in fired heater

There are two main heat transfer process – radiation and convection. Radiation occurs

within the radiation zone where fire from the burners heats up the tubes containing the

process fluid, in this case the reactants and products of the steam methane reform reaction. In

the convection zone, heat transfer is a combination of non-luminous and convective heat

transfer. The flue gas is the main medium for convective heat transfer to take place.

3.2 RADIATION ZONE DESIGN

Heat transfer to the radiant zone is the most important aspect of design for a fired

heater. An acceptable heat flux and metal tube temperature has to be achieved during design4.

3.2.1 Thermal Efficiency of Fired Heater

Heater efficiency is essential for determining the energy to be supplied through the

combustion process in the fired heater. It is the ratio of the amount of heat transferred to the

tubes to the amount of heat generated through combustion in the fired heater. The heater’s

efficiency is dependent on the following factors:

• Flue-gas stack temperature

• Excess air or oxygen

• Heat lost to the surrounding

• Design of the convection section in the fired heater

The flue gas stack temperature can be computed using the approach temperature,

which is the difference in the stack temperature to the inlet fuel temperature. Typical

approach temperature varies between 100-150°F1. Through HYSYS simulation, the stack gas

temperature is 565.6°C. The percentage heat available (thermal efficiency) can be derived

from the graph as shown below. Typically, heat efficiency can also be computed from the

following equation:

gasfuelofValueHeatingLower

etemperaturgasflueatavailableHeatEfficiencyHeater =



To obtain the furnace efficiency, a theoretical flame temperature has to be found.

However, a few assumptions have to be made to simplify calculations.

• Combustion of nitrogen is negligible • No carbon monoxide is formed

To calculate the heat released from combustion and the temperature of the products

formed, the enthalpy change of the combustion process can be considered2.

Fig 3.2.1a: Thermodynamic flow of combustion reaction

The total heat of combustion can be given as Heat of combustion = ∆HR + ∆HP + ∆H0C

Assuming adiabatic combustion, heat of combustion = 0 ∆HP = -∆HR - ∆H0C

Composition of fuel gas from PSA outlet consists mainly of CH4 and H2 where

number of moles of H2 is 3 times the number of moles of CH4. The other components will be

ignored for furnace efficiency computations as they are present in small quantities.

Hence, the main combustion reactions considered for calculations are

(1) CH4 + 2O2 → CO2 + 2H2O (2) 2H2 + O2 → 2H2O



Hence, Number of Moles of

CH4= 1; H2 = 3; O2 = 3.5; CO2 = 1; H2O = 5; N2 = (0.79/0.21) x 3.5 = 13.17

To compute ∆HR, assume

• Temperature of air is preheated to 150°C and

• Temperature of fuel feed from PSA unit is 40°C

• Flame temperature is 1900°C

• 1 mol of CH4 present

• Air comprise of 79% N2 and 21% O2

The Cp of the gases present at a flame temperature is obtained from literature.

Using the table above and with excel spreadsheet, iteration is performed to obtain the flame

temperature. A flame temperature is first assumed.

No. of moles % excess air Cp (KJ/mol-K)

Reactant

Fuel

CH4 1 68.05

H2 3 30.5

Air

O2 3.5 1.15 34.25

N2 13.16666667 1.15 32.39

Product

CO2 1 52.31

H2O 5 40.93

N2 13.16666667 1.15 32.39

O2 3.5 0.15 34.25

Table 2.1a: Excel spreadsheet used in calculation for flame temperature

∆HR = (68.05+3 x30.5) (50-25) + (34.25 x 3.5 x 1.15 + 13.17 x 1.15 x 32.39) (150-25) =

82525 KJ/mol

∆HP = (1 x 52.31 + 5 x 40.93 + 13.17 x 1.15 x 32.39 +3.5 x 0.15 x 34.25) x (flame T – 25) =

765.4 (flame T -25)

∆H0C can be computed as: ∆H0

C = 802800 + 241800 = 1044600 KJ/mol

Hence, 765.4 (flame T -25) = 1044600 – 82525 Flame T = 1389°C



The flame T is quite close to the original flame temperature calculated and hence flame T will

be taken to be 1390°C.

Fig 2.1: Flue gas profile of fired heater

Given the furnace profile: Furnace Effficiency = (Heat to process) / (Heat released by fuel)

Assuming stack temperature is 150°C,

Furnace efficiency = 765.4 (1390 -150) / (1044600 – 80930) = 0.9449

As it is usually not an adiabatic combustion process, heat is also lost to the

surrounding through the refractory walls. The value of heat loss is usually 2%3. Hence, the

overall thermal efficiency of the furnace is: 94.49 – 2 = 92.49%

Calculating the amount of heat to be supplied by the furnace, based on the energy

requirement specified by the SMR personnel,

Energy required = 81097.549.92

100××

= 6.45 x 108 KJ/h.

Assuming purged product from PSA contain a majority of methane gas for

combustion, the LHV of the fuel feed to the furnace will be approximately 50MJ/kg (5 x 104

KJ/kg). Hence amount of fuel feed needed is = 4

8

105

1082.5

×

× = 1.29 x 104 kg/h.



Given simulation from HYSYS, amount of PSA outlet is approximately 1 x 105 kg/h.

Since the CH4 content is not exactly 100%, despite the excess in quantity of fuel from PSA

outlet to that required, the totally energy that can be supplied will be the same as that of a

pure methane feed of lower quantity. Hence, the amount of fuel supplied from PSA outlet is

sufficient for supporting furnace combustion. However, in case insufficient fuel is supplied

from the PSA outlet due to equipment fault, a makeup fuel feed will be fed to the furnace.

This will be done through control instrumentation design. The amount of makeup feed will

then be 1.163 x 104 kg/h.

Fig 3.2.1d: Suggested

instrumentation control

for fuel gas inlet

control

3.2.2 Calculation for the number of reformer tubes

The number of reactor tubes within the furnace can be computed from an average heat

flux. Typical heat flux value for reformer unit is3 25000 BTU/h-ft2. With the number of tubes

computed, the mass velocity within the reformer tubes can then be computed. An excel

spreadsheet was used to compute the number of tubes from the heat flux value. The value of

heat flux is found in literature6 to be 25 000 BTU/hr-ft2.



Length of tube (ft) 39.2

Internal diameter of pipe (ft) 0.417

Surface area of tube (ft2) 51.3277

Total energy required in KJ/h as derived from HYSYS 5.97 x 108

Average flux in KJ/hr-ft2 as specified in literature 25000

No. of tubes 465

Table 3.2.2a: Computation for no. of tubes with heat flux

The number of tubes was calculated using the following equation:

)()(

literaturefluxheatAveragetubeoflengthpipeofdiameterInternal

HYSYSderivedasrequiredenergyTotal=

××π

Hence it can be concluded that the number of tubes to be used for the reactor is

approximately 465 tubes. However, since the SMR personnel have obtained good conversion

with 450 tubes, 450 tubes will be used for further design considerations.

3.2.3 Calculation for mass velocity in reformer tubes

tubepertioncrosstubesofno

tsreacofflowmasstotaltubereformerpervelocityMass

sec.

tan)(

×=

Total SMR feed load (kg/h) 182400

Total SMR feed load (lb/s) 111.4667

Internal diameter of pipe (inches) 5

Internal diameter of pipe (ft) 0.417

cross section of pipe in (ft2) 0.0137

mass flow velocity in lb/s (ft2) 18

No. of tubes 450

Table 3.2.3a: Computation for no. of tubes from mass velocity

The mass velocity of the fluid in the tube can be found to be around 18 lb/s ft2. The

minimum mass flow velocity required of 15 lb/s ft2 is satisfied.



Turndown consideration

It is important to consider turn-down and possible stream recycling. Usually,

turndown of 60%16 is taken into account. 60% x 15 = 10.8 lb/s ft2.

However, as the furnace supports a reaction process, recycling will not be considered.

3.2.4 Calculation of reformer tube thickness

HK40 (25% Cr, 20% Ni) are usually used for steam reformer tubes4. HK40 is chosen

for its high creep and rupture strengths. It is also resistant to hot gas corrosion and hence is

usually employed in steam methane reforming processes5. It provides creep resistance up to

980°C. This makes it suitable for the current design where the highest temperature 826.85°C6.

A standard code formula is employed to calculate the minimum wall thickness required7.

in

kPa

psikPa

ksi

psiksi

inm

inm

kPa

psikPa

PES

FCAd

P

ta

i

wall 399.0

1

1448.027006.085.0

1

10002.3

039.00254.0

1

2

127.0

1

1448.02700

6.0

2min =××+××

+××

×=

−

+

=

P = Max Pressure (to be in psig) = 2700 kPa (highest pressure at inlet);

di = Inner Diameter (to be in inches) = 0.127 m;

FCA7 = 10-year corrosion allowance (to be in inches) = 0.039 in;

Sa7 = Minimum Creep Stress for HK40 (to be in psi) = 3.2 ksi;

E 6 = Weld Efficiency Factor = 0.85.



3.2.5 Selection of material for reactor tube in radiation section

A series of procedure has been developed to aid the selection of material for reactors8.

1. Define the conditions of exposure (eg. Temperature and pressure)

2. Explore available materials

3. Identify the suitable material

4. Evaluate the material

In the steam methane reforming process, the reactor will be exposed to high

temperature of approximately 1130K and pressure of 2700kPa9. In addition, there are chances

of hot gas corrosion due to the high mass velocity. In view of these conditions, stainless steel

is a suitable material for the construction of the steam methane reformer tubes. With higher

carbon content, stainless steel offers greater creep resistance than other metals. With the

addition of nickel and chromium, resistance to carburization and creep is enhanced. There are

different grades of heat resistance steel, namely, HH, HK, HD and HF. HK has been

specifically found to be of great use in steam methane reforming due to their high creep and

rupture strength even up to 1150°C.

Most importantly, it

offers resistance to

hot gas corrosion.

The following figure

shows the superiority

of HK40 metal

compared to other

grades. It shows the

relative tensile

strength of the

different stainless

steel grades.

Fig 3.2.5a: Tensile strength



In addition, the tube life can be determined as follows.

Fig 3.2.5b: Tube life of HK40 tube

From the figure, at pressure of 2700kPa (391psi) and temperature of 1130K, the

furnace tube life can be found to be more than 20.3 years. This provides a confirmation that

the material is suitable for use in steam methane reforming processes.

The follow figure shows the minimum stress to rupture for HK40 piping as compared

to other grades. The figure implies a lower performance of HK40 grade compared to HP

grades. However, since the maximum temperature is 1130K, which is low compared to the

maximum temperature at which these

stainless steel grades can withstand, the

pressure factor is taken for higher

consideration in selection of the suitable

metal type. Since HK40 is capable of

withstanding higher pressure, will be the

final choice of material for the reformer

tubes.

Fig 3.2.5c: Minimum stress Vs Temperature



3.2.6 Reformer inner tube diameter

Tube sizes ranging from 4 to 8 inches are usually used based on the standard nominal

pipe size. A 5 inch inner tube diameter has been chosen based on the tube thickness

calculated (0.399 inches) and the standard nominal tube sizes1. Hence a Schedule 8016 tube

constructed from HK40 stainless steel will be used for the reformer tubes.

3.2.7 Furnace layout and design

3.2.7.1 Side Fired Heater

A side fired heater with vertical tubes has been used for simulation of the SMR

reaction. Hence, a side fired heater design will be proposed for the furnace type.

Side fired furnace has a few advantages. It allows the adjustment and control of the

tube wall temperature. The maximum temperature will be at the outlet of the reformer tube

while the highest heat flux is at a relatively low temperature. The side fired furnace offers

more flexibility in design and operation10. Side fired configuration also allows a counter-

current flow of flue gas and process fluid which yields a higher heater efficiency.

A typical side fired heater has the following configuration as found

in literature10.

However, given the large number of tubes, it is not economically feasible to line the

tubes in two rows as a large amount of space will be needed. Hence, 4 rows of tubes will be

proposed, each row comprising 450 /4 = 112 or 113 tubes.



3.2.7.2 Distance between burners

The distance

between the burners

is kept at 4m. This is

to ensure a safe

distance between the

tubes and the two

burners. The 2-D

sketch is shown

below.

Fig 3.2.7.2a: Proposed side-fired heater design (radiation + convection zone)

Given the tube dimensions as computed and that tube pitch is taken as twice the tube

outer diameter, and taking the allowance from the refractory wall to be 1 m in total,

Tube dimensions

Length (m) 11.95

I.D (m) 0.127

O.D (m) 0.147

Number of tubes per row 113

Tube pitch (where D =

outer diameter) 2 x O.D

Allowance

tubeperdiameterOuterpitchTuberowspertubesofNo

furnaceofLength

+

××= .

Hence length of furnace is approximately 33m.



The pigtails will be approximated to be 0.5m each. Pigtails are used to allow

expansion and contraction of tubes during start up. It minimises the need for joints and

welding that may fail when there is too much stress. Inlet pigtails are silicon killed while

outlet pigtail is made up of high alloy material.

Height of the radiant section will be taken to be: 11.95m + length of pigtails = 13m.

A 3-dimensional proposed design is as shown below.

Fig 3.2.7.2a: 3-D view radiation zone of proposed side-fired heater

3.2.7.3 Burners used at Side Walls

Premix burners will be used for the side wall. This is because they offer better

linearity, where excess air remains more nearly constant at turndown. Air will be drawn in

through the primary box register and mixed with the fuel before it flows to the furnace

firebox. Good mixing has to be ensured so that a short non-yellow flame can be obtained.

This is to prevent the flame from being in contact with the reformer tubes and cause locus

increase of temperature on the reformer tube. Long flames cause tube failure in the long run

and soot blower may be necessary to clean the heating surface. The figure below shows a

typical premix burner11.



Fig 3.2.7.3a: Pre-mix burner

However, usually gaseous fuels provide non-luminous flames4.

3.2.7.4 Determination of number of burners

The length of the furnace box is given to be 33m. For maximum heat distribution, the

centre to centre distance between burners should be 1m. Hence there would be approximately

32 burners along the length of the furnace. Since the height of the furnace is 13m, the number

of burners along the height of the furnace is 12. The layout on the refractory wall is shown

below. The total number of burners used will be 3072 burners.

Fig 2.7.4a: Side-fired heater burner arrangement



3.2.8 Computations for flue gas temperature

The flue gas temperature as obtained from HYSYS is 565.6°C. However, the

conversion reactor is used in the HYSYS design, which is a steady state module. Hence, more

calculations should be done to clarify the flue gas temperature. These set of calculations take

into account the dynamic state of the furnace.

3.2.8.1 Cold plane area

The cold plane area, which is the projected area of reformer tubes, is calculated as

follows13:

Acp= exposed tube length x centre to centre spacing x number of tubes excluding shield tubes.

= 11.95 x 2 (pitch) x 0.147 (outer diameter of pipe) x 450 = 1584 m2.

3.2.8.2 Refractory area

The refractory area is defined as the inside surface of the shell minus the cold plane

area. The equation for computation of the refractory area is as follows:

Aw = 2[W(H+L) + H x L)] = 2[16(13+33)+33 x 13] - Acp = 746 m2.

3.2.8.3 Absorptivity, α

α = 1- [0.0277 + 0.927 (x -1)] (x-1) ; where x refers to the pitch. Since pitch is 2, α = 0.879.

3.2.8.4 Sum of product of area and the absorptivities in the radiant zone

The equation for calculation is shown below: αAR = αAshield+ αAcp

Assuming the Ashield is negligible, then αAR = αAshield + αAcp

αAR = αAcp AR = Acp

Hence AR = 1584m2.



3.2.8.5 Mean beam length

L = (2/3)(furnace volume)1/3 = 12.7m

3.2.8.6 Partial pressure of CO2 and H2O

The main combustion products are CO2 and H2O.

P = 0.288 – 0.229x + 0.0090x2; where x is the fraction of excess air taken to be 0.15.

Hence P = 0.256.

3.2.8.7 Product of partial pressure and mean beam length

PL = 0.256 x 12.7 = 3.24

3.2.8.8 Mean refractory tube wall temperature

Tt = 100 + 0.5 (T1 + T2)

From the SMR personnel: T1 = 539.4°C and T2 = 862.85°C.

Hence Tt = 783°C = 1384°F

3.2.8.9 Two main equations that will be used for iteration to find Tg (flue gas temp)

3.2.8.9.1 Radiant zone heat transfer

( )tg

tg

R

R TTTT

FA

Q−+

+−

+= 7

1000

460

1000

4601730

44

α

3.2.8.9.2 Radiant zone heat balance

−−++=

n

g

n

L

n

f

n

a

R

n

R

R

Q

Q

Q

Q

Q

Q

Q

Q

FA

Q

FA

Q1

αα

The unknowns in the equations also require approximation of Tg to be made.



n

f

Q

Qrefers to the enthalpy of the fuel feed and is taken to be negligible as it is not preheated.

QR is related to Qn by the efficiency.

Qn = QR/efficiency where efficiency = 94% as computed earlier

n

a

Q

Q refers to the enthalpy of the preheated air and will be taken from HYSYS simulation.

3.2.8.9.3 Enthalpy of the flue gas as a function of Tg (flue gas temp)

−

−+= 1.01000

1.01000

TTba

Q

Q

n

g

Z = fraction excess air

a= 0.22048-0.35027*z+0.92344*(z)^2; b=0.016086+0.29393*z-0.48139*(z^2)

3.2.8.9.4 Emissitivity of the gas Ф

Ф = a + b(PL) + c(PL)2 where PL was calculated earlier on

Z = (Tg+460)/1000

a= 0.47916-0.1984*z+0.022569*(z^2); b= 0.047029+0.0699*z-0.01528*(z^2)

c= -0.000803-0.00726*z+0.001597*(z^2)

3.2.8.9.5 Exchange factor F

F = a + b Ф + c Ф2 Z = Aw/αAR

a=0.00064+0.0591*z+0.00101*(z^2); b=1.0256+0.4908*z-0.058*(z^2)

c=-0.144-0.552*z+0.04*(z^2)



After performing iteration using goal-seek tool in excel spreadsheet, the flue gas

temperature derived is 417°C. However the temperature as obtained from HYSYS is

565.6°C. However, since the heat exchanger network person-in-charge has decided to use

565.6°C for calculations in stream-matching, further computations and design for convection

section will make use of this value.

Since the temperature is low, a check is

carried out to ensure that the dew point of the

flue gas is not reached. The graph below shows

the dew point temperature of flue gas at different

temperatures.13.

Fig 3.2.8.9.5a14: Dew point of flue gases versus fuel sulphur

Given that the excess air is 15%, and that there is 0wt% sulphur in fuel, the dew point

if about 130°F, which is lower than the flue, gas temperature computed (410°C). Hence the

flue gas temperature computed is reasonable.

3.2.9 Residence Time

Residence time = volume of each reformer tube / volumetric flow rate of reactant gas

= π x (D/2)2 / (mass flow per tube/ density)

= 1.286s



3.3 CONVECTION SECTION

The convection section is used to preheat streams from other parts of the plant. This

helps to maximise the amount of energy that can be extracted from the furnace. Shield tubes

are omitted as convection tubes are not receiving direct heat from the flame from the

proposed design of the fired heaters. As discussed with the Heat Exchanger Network person-

in-charge, three process streams will have to be heated and the heating scheme will be as

shown below.

Steam will first be generated followed by heating up the SMR feed and finally

preheating the air fuel feed.

3.3.1 Convection design – Finned tubes

In the design of finned tubes, the following equation will be used.

)(LMTCDU

QAc

c

c= ; where ( )[ ])/(ln

)()(

011

011

LsLg

LsLg

TTTT

TTTTLMTD

−−

−−−= ;

T L0 and TL1 = inlet and outlet temperature of process fluid (respectively)

Tg1 and Ts = temperature of incoming and outgoing flue gas (respectively)



From the Heat Exchanger Person-in-charge, the following information is derived.

Process streams Heat duty KW LMTD Uc (W/m2K)

Steam 32090.7 87.5,360.7 10

Natural Gas 7307.5 61.4, 334.6 6.68

Air 4720.5 119.5,392.7 8.78

The total area is then calculated to be 8858.7 + 3268.6 + 1368.1 = 13495.4m2.

From the HEN person-in-charge, the desired outer diameter is 0.01905m = 0.75 in.

From the vendor of finned tubes (Vulcan Tubes), an appropriate fin tube is chosen.

The fin dimensions17 are shown below:

Number of fins per inches 7

Fin thickness (in.) 0.06

Fin height (in.) 0.625 = 0.015m

Surface area (sq ft per linear foot) 3.39

Total length of tubes needed = 145251 / 3.39 = 44015ft = 13338m

Section Total Length of each section (m)

Steam 8755

SMR feed 3126

Air 1352

Since the length of each tubes is very long, pressure drop will be high and hence the stream

has to be split into different tubes to prevent high pressure drop.

Section Number of tubes

steam 1250

SMR feed 446.47

Air 193

Number of tubes required = 13338 / 7 = 1905 tubes



To ensure a mass flow of flue gas to be 1.7kg/s-m2, the

number of tubes is computed as:

Flow rate of flue gas is 86.88kg/s;

Cross section area of tubes = 86.88/1.7 = 51m2.

Width of cross section = 51/7 = 7.29m

Number of tubes along the cross section

= 7.29 / (2 x (0.01905+0.015=0.03405)) = 107 tubes

Hence the numbers of rows of tubes are 404 / 22 = 17.8 rows

Assuming the same pitch, height of convection section = 18 x 2 x 0.03405 =1.21m

Final dimension of the convection box is: 1.21m (Height) × 7m (Length) × 7.29m (Width)

3.3.2 Design parameters for convection tubes

Dimensions and tube material as provided by Heat Exchanger Network (HEN) counterpart:

Thickness (m) 0.002

Outer diameter (m) 0.01483

Tube nominal size15 Schedule 10

Tube material Carbon Steel

The minimal thickness to withstand the creep of carbon steel is found using the

equation as used for HK40 calculated above (for thickness of reformer tube minimum wall

thickness), PES

FCAd

P

ta

i

wall6.0

2min

−

+

=

Given that the creep rupture strength of carbon steel is 54000 psi, the thickness is

0.00285in. Hence the minimum requirement is satisfied.



3.3.3 Pressure drop in the tubes present in furnace

The typical pressure drop for crude unit heaters is between 150-250 psi. An additional

20-25 psi is added for fouled tubes. In order to satisfy this condition, the number of passes

and tube size has to be optimized.

Flow of the fluid within the furnace pipes is turbulent, via Reynolds number: µρVDi=Re .

Thus, a correlation developed by Haaland19 was used to determine the Fanning friction factor:

( ) ( )[ ]9/10

10 7.3/Re/9.6log6.3/1 if Def +−=

The following conditions must be satisfied to accurately determine the friction factor with

this correlation: (a) 48 104Re10 ×≥≥ ; (b) 0/05.0 ≥≥ iDe

Otherwise ff = 16/Re (for laminar flow)

Computation of the frictional head loss for a straight pipe is evaluated using the relation:

)2(/2 2

ifL gDLVfh =

The pressure drop across the straight pipe is then given by: Lp hP ρ=∆

Presence of 180° bends within the 2 sections also contribute to the pressure drop

because the direction of flow changes. For each bend, a friction loss factor of K=1.6 is used

to compute the head loss. Subsequently, the pressure drop is obtained:

)2(/2 gKVhP LB ==∆ ρ

The total pressure drop in the tubes is evaluated by addition of the pressure drop

across straight pipes and the bends: 21 PPPT ∆+∆=∆



The pressure drop for the following streams is computed. For the SMR stream,

pressure drop through bends is neglected as the tubes are straight.

Stream SMR tubing

in radiation

section

Steam in

convection

section

SMR feed in

convection

tubes

Preheated air

in convection

tubes

Re 852 7934 7334 230.67

Fanning friction factor

(assuming smooth tube)

0.0188 0.00202 0.00218 0.0694

Friction head loss 0.103 0.142 346.6 154062

Pdrop through straight

tube (psi)

- 112 3982 180247

Pdrop through bends

(psi)

- - - -

Total Pdrop (psi) 0.109 112 3982 180247

It is noted that the pressure drop across the tube for preheating SMR feed and for

preheating air feed is much higher than the typical value. However, as this design is based on

the inner diameter as supplied by the heat exchanger network person-in-charge, this problem

will only be brought up for further mitigation on the best diameter for the convection tubes.

The pressure drop for SMR tubing is negligible, which is ideal for the steam methane

reforming reaction.



3.4 STACK DESIGN

The stack is designed to direct the flue gas out of the furnace into the atmosphere as

well as to achieve a draft of required combustion air through the furnace. The stack height

must be sufficient to achieve this flow without imposing a positive pressure on any part of the

furnace chamber18. The usual practice is to maintain a small negative pressure in the furnace

to enable the introduction of air from the atmosphere. It also allows for the removal of

undesirable products from the furnace. The required stack height is dependent on the

temperature of the flue gas leaving the convection section and the difference in density of the

flue gas and the atmospheric air.

3.4.1 Stack diameter

An acceptable velocity for the flue gas velocity is found12 to be 7.6m/s. Assuming that

the stack is a uniform cylinder,

Diameter = [(volumetric flow rate of flue gas) / (π x flow rate)] ^ 0.5 = 4.96m

3.4.2 Pressure Drop across stack

3.4.2.1 Stack exit loss

The stack exit loss is computed as follows: )273(/176.0 2

1 +=∆ ag TKVP

Velocity of flue gas = 11.5m/s ∆P1= 0.0783kPa

3.4.2.2 Frictional Loss in stacks and ducts

The flow in the stack is turbulent and hence the von Karman’s equation is used.

[ ] 28.2/log4/1 10 += eDf sf )(/2 2

2 gDHVfP sgf=∆

Assume the roughness factor is 0.5, ff= 0.86. H is taken to be 4m as an initial guide.

∆P2= 0.0186 kPa



3.4.2.3 Stack entrance loss

Stack entrance loss takes into account the full velocity head loss due to a change in

direction as the stack gas exits the furnace.

)273(/176.0 2

3 +=∆ stackg TVP ∆P3= 5.5×10-5 kPa

3.4.2.4 Flue gas pressure drop through the convection section

Gunter Shaw’s correlation is used for pressure drop of a bank of helical bank tubes of

staggered arrangements.

6.04.0

10

2

41022.5)/(

×=∆

T

L

T

ev

sevwg

p

S

S

S

d

d

LfGP

φρρ ∆P4 = 1.11×10-5 kPa

3.4.2.5 Pressure drop at the top of the radiant section

A vacuum of 2 mm H2O gauge just below the convection section is to be maintained

to prevent leakage of flue gas through the casing of the furnace. Hence ∆P5 = 0.0020kPa.

3.4.2.6 Pressure gain at the convection section

The stack effect at the convection section brings about a pressure gain in the furnace.

This gain is caused by the density difference between the hot flue gas and the ambient air

outside.

∆P6= 2.7 x 10-2 kPa ;

Total Pdrop across stack = ∆P1 + ∆P2 +∆P3 + ∆P4 +∆P5 +∆P6 = 0.0720kPa



3.4.3 Stack Height

A stack height of 4m has been approximated

for calculation of pressure drop through the stack.

The stack height is calculated again to ensure that the

approximation is correct. According to the Code of

Practice on Pollution Control by National

Environment Agency (NEA), the stack height should

be at least 15m from the ground.

This is so that the hot stack gases are

discharged at a safe height with respect to the

surrounding equipment in the plant. In addition, the

flue gas may contain pollutants such as SOx, NOx and

particulates. Hence, the stack must be designed to

discharge these gases in a manner that avoids causing

a local pollution problem.

The equation used to calculate the stack height:

−=gaa

atmgd TTPHP 1135.0

Pd = 0.0720kPa. H is found to be 10m. Hence the average height is taken to be 8m.



3.5 MATERIALS FOR CONSTRUCTION OF FURNACE BODY & ADDITIONAL AUXILIARIES

3.5.1 Refractory walls

The refractory walls are to be made of strong material that can withstand high

temperature. It also must resist abrasion and flue gas and most importantly, it should have a

high insulation to prevent heat loss to the surrounding.

In this design, silica brick (97-98% silica) with a thickness of 5-8 inches is selected to

line the furnace walls. It has the ability to retain its strength at high temperatures. A highly

porous fire clay insulating firebrick (1”) is placed between this lining & the metal casing.

The silicon carbide coating is light, low in thermal conductivity and sufficiently

resistant to temperature for the use on the hot side of the furnace wall. Thus, it permits thin

walls of low thermal conductivity and low heat content. The low heat content is particularly

valuable in saving fuel and time on heating time.

The properties of the silicon carbide and insulating wall are shown below.

Properties Silicon Carbide Insulating Brick

Thermal shock resistant Excellent Excellent

Hot strength/

Deformation under hot loading/

Permeability

Excellent/

Excellent/

Very Low

Poor/

Poor/

High

Fusion pt (oF) 4175 Varies

Bulk density lb/ft3 160 30-75

Composition SiC 80-90% Varies

To further confirm that the refractory material chosen will be able to withstand high

temperature from the flames, Stefan-Boltzman equation will be used: 4Tqr σ=

Given that the radiant heat flux is about 25,000 BTU/h-ft2, which is 7.9 × 104 W/m2, T (wall

temperature) = 1086K = 1495F which is less that the fusion pt of silicon carbide and hence

this material is suitable for use.



3.5.2 Stack Walls

Stainless steel will be used15 for construction of stack walls instead of insulating brick.

This reduces the cost and will exert less stress to lesser weight of the material. Stainless steel

melts at around 1370°C. Since the maximum temperature within the stack is approximately

565°C, stainless steel is suitable to form the stack wall.

However, it is important to note that since metal is involved, the temperature within

the stack should be kept above 150°C, which is above the dew point of water to prevent

condensation and thereby the formation of acid which will corrode the metal.

3.5.3 Additional auxiliaries

3.5.3.1 Air Preheaters

There are commercially available air preheaters to heat up the furnace air feed. One of

the commercially available air preheater is the Rekuluvo® Recuperative Air Preheater. Air is

preheated prior to burning in the furnace to ensure higher heat recovery. This

The good accessibility to heating surfaces allows easy maintenance. In addition, it is

corrosion resistant and does not have any mechanical moving parts that need additional

power supply.



3.5.3.2 Forced Draft Fan

Forced draft has to be installed at the inlet of the furnace to draw air supply to the

furnace. This is because after being preheated by heat exchangers, the pressure of the air

supply drop by 3 psi for each heat exchanger. Having passed through 3 heat exchangers, the

pressure drop would be 9 psi in total.

The final pressure before entering the fired heater might be 14.5 – 9 = 5.5 psi, which

is very low. The forced draft will be used to increase the pressure of the air supply to 1atm

prior to feeding into the fired heater.

3.5.3.3 Induced Draft Fan

Induced draft fan is placed at the outlet of the furnace to draw the flue gas out of the

stack. A pressure of 2mmH2O less than atmospheric pressure is maintained. The proposed

force and induced draft fan are shown below.

Since both types of draft are used, the set-up is known as balanced draft. The fans

will be chosen in a way that the pressure is slightly below atmospheric pressure. This ensures

safe operation and reduces leakage of air into the furnace.



3.6 COST ANALYSIS

3.6.1 Purchased Equipment Costs

Preliminary cost estimation was done to estimate the furnace and its auxiliary

equipment cost using the CAPCOST program developed by Turton.

3.6.1.1 Costing for Furnace

The bare module cost of reformer furnace before accounting for inflation is calculated

from Equation (3-1). Equation (3-2) gives the pressure factor (Fp) for the furnace. As carbon

steel is the base material used, the material factor, FM, is 2.1

.

(3-1)

for P < 10 barg (3-2)

(3-3)

where Ft (superheat correction factor for steam boilers) = 1 for heaters and furnaces

Identification number for HK40 alloy steel is 54, hence bare module factor FBM = 2.5

The various parameters that will be used for cost estimation:

Parameter Unit Value

A, Heat Duty KW 1.65 x 105

Pbarg barg 1

FM 2.5

FP 1

CBM= USD 1.545 x 1012

As the data for the equations were obtained during May to September 2002 when

Chemical Engineering Plant Cost Index (CEPCI) was 395.6, inflation should be accounted

for using the CEPCI of 595.1 in the last quarter of 2007 to USD 2.32 x 1012.



3.6.1.2 Costing for Air Preheater

The bare module cost of the air preheater was approximated to that of a flat plate heat

exchanger because of the difficulty of getting the actual cost from vendors. Carbon steel was

chosen as the base material. Similarly, the bare module cost is calculated with equation 3-4.

)21.196.0( pM

o

pBM FFCC += (3-4)

Parameter Unit Value

A, Area M2 160

Pbarg barg 1

FM 1

FP 1

After taking inflation into account by using the CEPCI of 512 in the last quarter of

2006, the estimated cost of the air preheater is US$256,000.

3.6.1.3 Costing for Induced Draft Fan and Forced Draft Fan for Air Preheating System

The induced and forced draft fans selected are centrifugal fans, thus the bare module

cost can be approximated with that of the centrifugal radial fan. For both, carbon steel was

selected as the bare material since only flue gas and combustion air will be in contact with

them at relatively low temperatures.

pM

o

pBM FFCC = (3-5)

The table below shows the Fans Bare Module cost parameters

Using the same CAPCOST Program, and after taking inflation into account, the

estimated cost of the induced draft fan is US$13,500 and forced draft fan is US$52,000.



3.6.1.4 Burners

There are 3072 small premix burners used for this operation. Assuming each of the

premix burner cost USD 100, the total cost is USD 307200.

3.6.2 Utility Cost

Electricity (440V, 3-phase, 50Hz) – USD 100/MWh

Natural gas feed is not considered for furnace operation as the off gas from the PSA is

sufficient to supply enough heat energy required.

3.6.2.1 Electricity cost

The units of the furnace that runs on electricity are the induced and forced draft fans.

Hence, the electricity consumption will be based on the Horsepower rating of the individual

models that were selected according to the required capacity of volumetric gas flow rate.

Fan Horsepower Power (KW)

Induced Draft 25 19

Forced Draft 30 22

Given the power rating, the amount of electricity to operate both fans is US$14,300/year.

3.6.3 Total Annual cost

Assuming 15 years of operation:

Total annual cost = total bare module cost / 15 years + operating cost per year

Total annual cost = USD 1.55 x 105 million per year



3.7 SUMMARY & CONCLUSION

In this project, a fired heater design for the steam methane reforming process has been

developed. The design methodology was presented. Design of the fired heater began with

consideration of the reaction heat energy requirement. The radiation section is then designed

with heuristics and maximum temperature specifications in mind.

After the completion of the radiation zone design, the convection design was explored

to increase the efficiency of the fired heater. The streams to be heated in the convection zone

were identified and the stream data and tube dimensions were obtained from the heat

exchanger person-in-charge. With the information available, the finned tube arrangements

were determined and the sizing of the convection section was obtained.

Both design process paid attention to heuristics and chances for optimisation.

Finally, stack design was carried out to meet specifications by governmental bodies.

Costing was then performed to determine the total annual cost of the fired heater constructed.



3.8 SPECIFICATION OF FIRED-HEATER

Service: Steam Methane Reforming

Design Duty: 5.6 x 108 BTU/h Unit: Vacuum Unit heater

No. of heaters 1 Type: Box

Design Radiant Section Convection Section (Total) Service SMR Preheating Streams

Heat absorption (mmBTU/h)

561 108.6 15.9 24.7

Fluid SMR Process Fluid

Superheated steam

Air SMR feed

Allowable pressure drop (psi)

150-200

Allowable average heat flux (BTU/h-ft2)

25000

Fouling factor 0

Residence time N/A

Inlet Conditions Temperature (°C) 539.4 253.3 25 25

Pressure (kPa) 2679 4200 100 4000

Liquid flow (kg/h) N/A 1.743 x 105 N/A N/A

Vapour flow (kg/h) 1.835 x 105 N/A 2.997 x 105 4.3 x 104

Liquid density (kg/m3) N/A 790.7 N/A N/A

Vapour density(kg/m3) 7.059 N/A 1.167 28.99

Viscosity(cST) 3.497 0.1336 16.12 0.4180

Specific heats (KJ/KJmole-C)

46.29 99.98 29.24 40.72

Thermal conductivity (W/m-K)

0.08 0.6121 0.02586 0.0369

Design Radiant Section Convection Section (Total)

Outlet Temperature (°C) 851.9 254.3 80.30 250

Pressure (kPa) 2630 4179 79.32 3979

Liquid flow (kg/h) N/A N/A N/A N/A

Vapour flow (kg/h) 1.835 x 105 1.743 x 105 2.997 x 105 4.3 x 105

Liquid density (kg/m3) N/A N/A N/A N/A

Vapour density(kg/m3) 3.55 20.27 0.7804 15.10

Viscosity(cST) 8.662 0.8705 27.33 1.192

Specific heats (KJ/KJmole-C)

38.47 43.58 29.57 49.05

Thermal conductivity (W/m-K)

0.2003 0.05 0.02972 0.07



3.9 REFERENCES

1. Furnace/ Fired Heater Design and Control Lecture notes

2. Robin Smith. (2005). Chemical Process Design and Integration. John Wiley & Sons.

3. W. L. Nelson. (1985). Petroleum Refinery Engineering, Auckland. McGraw-Hill

4. V. Ganapathy. (1982). Applied heat transfer. PennWell.

5. Role of Alloying Elements. Retrieved from World Wide Web on 16 Mar 2008:

http://www.valve-world.net/pdf/11022.pdf

6. Retrieved from World Wide Web on 30 Mar 2008:

http://www.fitness4service.com/publications/pdf_downloads/Jaske-Shannon%20Paper.PDF

7. Inspection and Remaining Life Evaluation of Process Plant Equipment. Retrieved from

World Wide Web on 16 Mar 2008: http://www.kubotametal.com/alloys/heat_resistant/HK-

40.pdf

8. Dilton, C.P. (1992). Materials selection for the chemical process industries.

9. Role of Alloying Elements. Retrieved from World Wide Web on 16 Mar 2008:

http://www.valve-world.net/pdf/11022.pdf

10. Rostrup-Nielsen, J. (1993). Steam Reforming Opportunities and Limits of the Technology,

Catalysis Today, Vol. 18, P305-324.

11. James R. Cooper, W. Roy Penney, James R. Fair. (2005). Chemical Process Equipment,

Second Edition: Selection and Design. Elsevier.

12. R.K. Sinnott, Coulson & Richardson's chemical engineering - Volume 6: Chemical

engineering design, Elsevier Butterworth-Heinemann (2005)

13. D.S.J. Jones. (1996). Elements of Chemical Process Engineering. John Wiley & Sons.

14. Melting Point of Iron – Jefferson Lab. Retrieved from World Wide Web on 20 Mar 2008:

http://72.14.235.104/search?q=cache:3kYX9gKVDVEJ:education.jlab.org/qa/meltingpoint_0

1.html+melting+point+of+steel&hl=en&ct=clnk&cd=1&gl=sg

15. James R. Welty, Charles E. Wicks, Robert E. Wilson & Geogory Rorrer. (2001).

Fundamentals of Momentum, Heat, and Mass Transfer – 4th Edition. John Wiley & Sons, Inc.

16. S. Singh, S Goyal. (2002). Fired Heaters in Chemical Process Industries CPECNews:P2-6

17. Retrieved from World Wide Web on 30 Mar 2008: http://www.vulcanfinnedtubes.com/

18. R.K. Sinnott. (2005). Coulson & Richardson's Chemical Engineering - Vol 6: Chemical

engineering design. Elsevier Butterworth-Heinemann.

19. S. E. Haaland. (1983). Trans. ASME, JFE: Vol. 105, P89

CN 4120: Design II Team 32: Ong Song Kun (U046829M) HTS Unit Design Report


Chapter 4 : HIGH TEMPERATURE SHIFT REACTOR

4.1 INTRODUCTION

Hydrogen is an extremely important compound required in our lives. It is used to

provide food, fuel and chemical resources for us. The largest processes using hydrogen are

ammonia synthesis, methanol synthesis, and hydrogenation. Without it, it would cause us to

live very differently in the world today. The production of hydrogen can be carried out using

steam reforming of any hydrocarbon source such as coal, methane, petroleum naphtha or

biomass. Methane is usually used due to its cheaper costs than the other hydrocarbon sources.

4.1.1 Water gas shift

During the steam reforming stage, side reactions happening in the steam reformer

would cause carbon monoxide to be formed, and this limits the production of hydrogen. Thus,

water gas shift reaction was developed to obtain more hydrogen from carbon monoxide.

This is the water gas shift reaction:

222 HCOOHCO +⇔+

which involves the reaction of carbon monoxide and water in the presence of a suitable

catalyst to form carbon dioxide and hydrogen.

There are three alternatives for carrying out the reduction of CO. [1]

1. Remove part of CO with iron catalyst in one bed. Then absorb CO2 and go to a second

bed of the same catalyst with a more favourable equilibrium since the product CO2 is

absent.

2. Conduct the entire reaction in a single bed on copper-zinc catalyst.

3. Remove part of the CO in a bed with iron catalyst and complete the removal in a

second bed of the more expensive copper-zinc catalyst.

The second and third alternatives are more attractive as the additional absorption equipment

in the first alternative creates added maintenance problems, particularly due to the corrosive

character of monoethanolamine, which is the usual absorbent used.



For the purpose of this design, I have decided to use the third alternative. This is because the

second alternative would be much more expensive. The cost of copper-zinc catalyst is 3 times

the cost of iron catalyst. Therefore, it is justified to use alternative 3, to achieve the required

CO reduction.

4.1.2 High temperature shift

222 HCOOHCO +⇔+ molkJH rxn /447.44−=∆

As seen from the enthalpy of the water gas shift reaction, it is an exothermic reaction.

Thermodynamically, the conversion of the reaction is favored at low temperature. The lower

the temperature is, the higher the conversion will be. However, at low temperatures, the rate

of reaction is slow. Though conversion is high, it might take a very long time for it to reach

that conversion equilibrium. Therefore, to ensure a high rate of reaction and a high overall

conversion, it is necessary to use a High Temperature Shift (HTS) followed by a Low

Temperature Shift (LTS). This mechanism is needed so that in the HTS reactor, the reaction

occurs at a reasonably high rate. Then the reaction is completed in the LTS, which would

ensure a reasonable overall conversion.

The higher temperature in the HTS reactor also allows recovery of the heat of reaction at a

sufficient temperature level to generate high pressure steam.

The HTS is usually conducted at a range of 315oC-480 oC. [1]



4.2 PROBLEM DESCRIPTION

After Steam Methane Reforming (SMR), the products would be thrown into the HTS reactor.

The objective of the HTS reactor is to reduce carbon monoxide composition to 3% (dry basis).

This was justified by literature data. [1] After which, the products would be thrown into the

LTS reactor for further conversion so that the exit carbon monoxide composition would be

0.7% (dry basis).

Feed specification

The relevant data of the outlet stream from SMR are as follows:

Table 4.2.1: SMR outlet stream data

Flow rate (kmol/hr) 14470

Pressure (kPa) 2610

Temperature (K) 1125

Table 4.2.2: SMR outlet stream composition

Component Mol fraction Molar flow (kmol/hr)

CH4 0.03493 505.4

H2O 0.34518 4994.8

CO 0.08737 1264.3

CO2 0.05451 788.8

H2 0.47513 6875.2

N2 0.00072 10.4

C2H6 0.00216 31.2

Steam-to-CO Ratio

According to literature, steam to carbon monoxide ratio must surely be more than 4:1.[1] The

optimum amount of steam to be used is based on economic considerations, such as the cost of

steam. Furthermore, using more steam requires equipment with a larger diameter due to a

greater flow rate.



From the above table, the steam to carbon monoxide molar ratio is 4:1. However, it was still

unable to achieve a reduction of CO to 3% (dry basis) at this steam ratio. Therefore, steam

was added to make the steam to carbon monoxide molar ratio 5:1. This was justified by

literature data, as they also used a steam to carbon monoxide ratio of 5:1. [1] Doing this

would reduce the amount of catalyst needed, as well as make it possible to achieve a

reduction of CO to 3% (dry basis).

Amount of steam added = ]8.4994)53.1264[( −× kmol/hr

= 1326.6 kmol/hr

After adding 1326.6kmol/hr of steam to make the steam to carbon monoxide molar ratio (5:1),

also cooling the inlet stream down to 627 K, and assuming a pressure drop of 20.88kPa

across the heat exchanger, the HTS inlet stream data are as follows:

Table 4.2.3: HTS inlet stream data

Flow rate (kmol/hr) 15796.6

Pressure (kPa) 2589.12

Temperature (K) 627

Table 4.2.4: HTS inlet stream composition

Component Mol fraction % composition

(dry basis)

Molar flow

(kmol/hr)

CH4 0.03200 5.33 505.4

H2O 0.40017 - 6321.4

CO 0.08004 13.34 1264.3

CO2 0.04993 8.32 788.8

H2 0.43523 72.56 6875.2

N2 0.00066 0.11 10.4

C2H6 0.00198 0.33 31.2



Product specification

Therefore, the problem is defined to reducing carbon monoxide from 13.34% (dry basis) to

3% (dry basis).

This is the table of the compositions of the HTS outlet stream after the single fixed bed

catalytic reactor was designed:

Table. 4.2.5 HTS outlet stream composition

Component Mol fraction % composition

(dry basis)

Molar flow

(kmol/hr)

CH4 0.03200 4.85 505.4

H2O 0.33994 - 5369.9

CO 0.01980 3.00 312.8

CO2 0.11017 16.69 1740.3

H2 0.49546 75.06 7826.6

N2 0.00066 0.10 10.4

C2H6 0.00198 0.30 31.2

Conversion of CO in designed HTS reactor = (1- 0.01980/0.08004) x 100%

= 75.27%



4.3 REACTION THERMODYNAMICS

4.3.1 Criteria for Chemical Reaction Equilibrium

The fundamental property relation for single-phase systems, provides an expression for the

total differential of the Gibbs energy:

d nG` a

= nV` a

dP @ nS` a

dT + Σiui dni

(4.3.1.1)

If changes in the mole numbers ni occur as the result of a single chemical reaction in a closed

system, then by substituting dni = v i dε , equation (2.3.1.1) gives:

d nG` a

= nV` a

dP @ nS` a

dT + Σiv i ui dε (4.3.1.2)

Because nG is a state function, the right side of this equation is an exact differential

expression; thus,

Σiv

iu

i=

∂ nG` a

∂εfffffffffffffffffffffF G

T,P

=∂G

t

∂εffffffffffffF G

T,P

Thus the quantity Σiv

iu

i represents the rate of change of total Gibbs energy of the system

with respect to the reaction coordinate at constant T and P. This quantity is zero at the

equilibrium state. A criterion of chemical-reaction equilibrium is therefore:

Σiv

iu

i= 0 (4.3.1.3)

The definition of the fugacity of a species in solution is as such:

µi= Γ

iT

` a+ RTln f

i

^

In addition, the following equation may be written for pure species i in its standard state at

the same temperature:

G i

o = Γi

T` a

+ RTln fi

o

The difference between these two equations is:

µi@G i

o = RTlnf

i

^

fi

o

ffffffff (4.3.1.4)

Combining equation (4.3.1.3) with equation (4.3.1.4) gives for the equilibrium state of a

chemical reaction: Σiv i G i

o + RTln fi

^

fi

o)f g v

i

H

J

I

K= 0



or Σi

v i G i

o + RTΣiln f

i

^

fi

o)f g v

i

= 0

or Πi

fi

^

fi

o)f g v

i

=@Σ

iv

iG i

o

RT

fffffffffffffffffffffffffff

where Πi

signifies the product over all species i. In exponential form, this equation becomes:

Πi

fi

^

fi

o)f g v

i

= K

4.3.2 Effects of Pressure on Reaction Equilibrium

The equilibrium state of a chemical reaction is given as:

Πi

fi

^

fi

o)f g v

i

= K (4.3.2.1)

Where Π represents the product over all species i, f is the fugacity of species i in solution,

fi

o is the fugacity of species i at standard state, and the equilibrium constant K is a function

of temperature only and is defined by:

K = exp@∆G

o

RT

fffffffffffffffffffff g

(4.3.2.1a)

The standard state for a gas is the ideal-gas state of the pure gas at the standard state pressure

Po of 1 bar. Because the fugacity of an ideal gas is equal to its pressure, fi

o = Po for each

species i. Thus for gas-phase reactions fi

^

fi

o)= f

i

^

Po) , and equation

(4.3.2.1) becomes:

Πi

fi

^

Po

ffffffff

h

j

i

k

vi

= K (4.3.2.2)

Equation (4.3.2.2) relates K to fugacities of the reacting species as they exist in the real

equilibrium mixture and these fugacities reflect the non-idealities of the equilibrium mixture.

The fugacity is related to the fugacity coefficient by:

fi

^

= Φi

^y

iP



Substituting this equation into (4.3.2.2) yields an equation which relates the pressure to the

composition:

Πi

yiΦ

i

^d e v

i

=P

Po

fffffffff g@v

K (4.3.2.3)

Where, v a P v i . Assuming that the equilibrium mixture is an ideal solution, then each Φi

^

becomes Φ i . Thus, equation (4.3.2.3) becomes:

Πi

yiΦ

i

b c vi

=P

Po

fffffffff g@v

K (4.3.2.4)

Each Φ i for a pure species can be calculated from a generalized correlation once the

equilibrium temperature and pressure is specified. For low pressures or high temperatures, the

equilibrium mixture behaves as an ideal gas where Φ i

^= 1. Thus, assuming that the

equilibrium mixture is an ideal gas, equation (4.3.2.4) reduces to:

Πi

yi

` avi =

P

Po

fffffffff g@v

K (4.3.2.5)

In the WGS shift reaction, the stoichiometric coefficients of the reactants and products are all

1 which means that v = P v i = 1 + 1 @1 @1` a

= 0. Therefore, equation (4.3.2.5) reduces to:

Πi

yi

` avi = K (4.3.2.6)

From equation (4.3.2.6), it can thus be seen that the equilibrium constant of the WGS reaction,

K is independent of pressure. Thus, the pressure conditions within the HTS reactor will not

affect the equilibrium of the reaction.

4.3.3 Effects of Temperature on Reaction Equilibrium

From the first law of thermodynamics for a closed system of n moles, is as such for the

special case of a reversible process:

d nU` a

= dQ + dW (4.3.3.1)

As applied to this process, dW =@Pd nV` a

and dQ = Td nS` a

. Combining these three

equations gives:

d nU` a

= Td nS` a

@Pd nV` a

(4.3.3.2)



The enthalpy and the Gibbs energy are defined as:

H a U + PV (4.3.3.3)

G a H @TS (4.3.3.4)

Upon multiplication by n, equation (4.3.3.3) may be differentiated to give:

d nH` a

= d nU` a

+ Pd nV` a

+ Vd nP` a

(4.3.3.5)

Substituting equation (4.3.3.2) into (4.3.3.5),

d nH` a

= Td nS` a

+ nV` a

dP (4.3.3.6)

In the same way, equation (4.3.3.4) may be multiplied by n and differentiated to give:

d nG` a

= d nH` a

@Td nS` a

@Sd nT` a

(4.3.3.7)

Equation (4.3.3.6) and equation (4.3.3.7) combine to yield:

d nG` a

= nV` a

dP @ nS` a

dT (4.3.3.8)

In the application of equation (4.3.3.8) to a one mole of homogeneous fluid of constant

composition, equation (4.3.3.8) simplifies to:

dG = VdP @SdT (4.3.3.9)

An alternative form of equation (4.3.3.9) which is a fundamental property relation that

follows from the mathematical identity is:

dG

RT

fffffffffff g

a1

RT

ffffffffffdG @

G

RT2

fffffffffffffdT (4.3.3.10)

Substituting equations (4.3.3.9) and (4.3.3.4) into (4.3.3.10):

dG

RT

fffffffffff g

aV

RT

ffffffffffdP @

H

RT2

fffffffffffffdT (4.3.3.11)

All terms in this equation are dimensionless. When applied in restricted forms,

@H

RT2

fffffffffffff =∂ G RT+

b c

∂T

ffffffffffffffffffffffffffffff

H

J

I

K

P

(4.3.3.12)

The relation between the standard heat of reaction and the standard Gibbs energy change of

reaction may be developed from equation (4.3.3.12) written for each species i in its standard

state:

H i

o = @RT2 d Gi

oRT*

b c

dT

ffffffffffffffffffffffffffffffffff

h

j

i

k (4.3.3.13)



By multiplying both sides with iν and summation over all species yields:

P viH

o = @RT2

d P vi

Gi

oRT*

b c

dT

ffffffffffffffffffffffffffffffffffffffffffffffffff

h

j

i

k (4.3.3.14)

By definition, ∆Goa P

i

viG i

o and ∆H

oa P

i

viH i

o. Thus equation (4.3.3.14) can be

expressed as:

∆Ho =@RT

2 d ∆Go

RT*b c

dT

fffffffffffffffffffffffffffffffffffff

h

j

i

k (4.3.3.15)

Substituting equation (4.3.2.1a), equation (4.3.3.15) becomes:

dln K

dT

ffffffffffffffffff=∆H

o

RT2

ffffffffffffff (4.3.3.16)

Equation (4.3.3.16) gives the effect of temperature on the equilibrium constant, and hence on

the equilibrium conversion. If ∆Hois negative, i.e. the reaction is exothermic, the equilibrium

constant decreases as the temperature increases. Conversely, K increases with T for an

endothermic reaction. Since the water-gas shift reaction is slightly exothermic with ∆H = -

41.1kJmol-1, thus the equilibrium constant increases with decreasing temperature. Thus, it is

desirable to operate at the lowest possible reactor inlet temperature to obtain maximum

removal of carbon monoxide.



4.4 REACTION KINETICS

The overall reaction is 222 HCOOHCO +⇔+

Using literature data [1], we have chosen Chromia-promoted iron oxide as our catalyst. This

catalyst has been used for many years for the shift reaction.

Rate Equation

The rate equation for this catalyst is shown below, and is assumed to represent midlife

activity:

b

HCOOHCO

CO

Kyyyykr

ρ

ψ

379

)/()( 222

−=−

Where,

k = rate constant )4900

95.15exp(T

−=

K = equilibrium constant )4578

33.4exp(T

+−=

(-rCO) = rate, lb moles CO converted / (lb catalyst) (hr)

T = temperature, K

yj = mole fraction of component indicated

ρb = catalyst bulk density, lb/cu ft

0.4=ψ for P > 20.0 atm

The manufacturer has subjected the rate equation to many tests, as well as observations on

full-scale plants. The rate constants are expressed on the basis of a reasonable “lined-out”

activity that the catalyst would maintain for a considerable time, if operating errors which

cause deactivation do not happen. The ψ term is the product of the total pressure (atm) and

ratio of the first-order constant at pressure P to that at atmospheric pressure and is a function

of pressure and Thiele modulus. Thus, it is considered that the effectiveness factor of the

catalyst has already been taken into account in the ψ term.



4.5 CATALYST

As mentioned, the catalyst used is chromia-promoted iron oxide. [1] The specifications of the

catalyst are as follow:

Table 4.5.1

Maximum operating temperature (oF) 890

Tablet size (inch) 0.25 x 0.25

Bulk density (lb/cu ft) 70

Particle density (lb/cu ft) 126

Catalyst poisons Inorganic salts, boron, oils, or phosphorous

compounds, liquid H2O is a temporary poison.

Sulfur compounds in an amount greater than

50ppm

Catalyst life 3 years and above, depends on care in startup

and operation (Use times up to 15 years have

been reported)



4.6 REACTOR

An adiabatic single fixed bed catalytic reactor is used for the HTS reactor.

4.6.1 Type of reactor

The reason why the reactor was chosen as a single bed was because the heat of reaction for

222 HCOOHCO +⇔+ was not that high. Therefore, temperatures in the reactor would not

rise by too much. A single bed would suffice to convert CO to its desired composition

without raising temperature too high such that conversion would be affected.

An adiabatic reactor was chosen because it is cheap and easy to maintain. It is not only the

lowest cost and simplest type of reactor, but its performance can be predicted reliably for

single phase reactions.

This is also justified by literature data as shown in the case study, where the author also used

an adiabatic single fixed bed catalytic reactor.

4.6.2 Reactor design

Vessel Design

Vessel costs are an important element in reactor design decisions. In the U.S.A. the American

Society of Mechanical Engineers has established a code for the design and fabrication of

pressure vessels. Similar organizations in Europe also have established codes. All such codes

give the minimum standards. Normally vessels as important as reactors are designed to

comply not only with a code but also with supplement specifications considered important for

a particular service. These can include special impact test requirements to assure against

brittle fracture, heat-treating specifications for steel in severe service such as high hydrogen

partial pressures. Below is a picture of the design of the vessel.



Fig. 4.6.2.1 Design of vessel

Corrosion Allowance

Although practices vary, on the average a material is selected that will not corrode more than

0.010 to 0.015 in. /yr. For a vessel of life 10 years, this approximates a corrosion allowance

of 1/8 in. Because of the many variables and unknowns associated with corrosion, a

minimum allowance of 1/8 in is specified for carbon steel and low alloy steel even if no

corrosion or erosion problems exist. For higher alloys, such as stainless steel, a lower

minimum of 1/32 in is often used.



Material Selection

In the case of hydrogen services which cause pitting, corrosion allowance may not be that

useful. Hydrogen destroys metal strength by producing cracks or blisters, but the thickness of

the metal is not reduced. At low temperature atomic hydrogen produce by thermal or catalytic

dissociation diffuses into the metal along imperfections, ultimately recombining to form

molecular hydrogen. The hydrogen pressure can increase to a point where it causes internal

and surface blistering. [4] At high temperatures, hydrogen diffuses even more rapidly and

forms methane by reacting with the carbon content of the steels. The larger methane molecule

builds up pressure that produces high internal pressure and ultimately cracks [5]. Neither of

these processes reduces the metal thickness. Thus one selects for high temperature service a

metal that will not be subjected to attack, containing a carbide stabilizing element such as

molydenum.

Thus the material ASTM A 387 Grade 22, Class 1 (2 ¼ Cr-1 Mo) was chosen for its

resistance to hydrogen attack.



4.7 METHODOLOGY AND CALCULATIONS

4.7.1 Weight of catalyst

Methodology

The following equations were used in the calculations for weight of catalyst used

Mass balance equation:

)()( COCO FrW ∆−=−∆ (Eq. 4.7.1.1)

Heat balance equation:

WHrTTcF TCOCOjjjpj ∆∆−−=−∑ + ))(()( 1, (Eq. 4.7.1.2)

))(( COCO HF ∆−∆−=

The heat capacities of gases were taken from literature text. [2]

Based on these equations and the rate equation, a MATLAB program was written based on

the following algorithm to find the mass of catalyst needed.

Algorithm

1. Input the inlet temperature of HTS in K.

2. Assume ∆W of 200lbs

3. Calculate (-rCO) at inlet conditions to increment, i.

4. Calculate (-rCO)avg = (-rCO)i + [(-rCO)i – (-rCO)i-1]/2 (skip for i=0).

5. Calculate new flow rates: Fi+1 = Fi ± (-rCO)∆W

6. 4. Calculate cp and (-∆HCO) @ Ti

7. Calculate ∆T from Eq. 4.7.1.2.

8. Ti+1 = Ti + ∆T

9. yi+1 = Fi+1/(FT)i

10. Mole fraction CO in dry gas = [yCO/(1-yH2O)]i+1

11. If mole fraction CO in dry gas is more than 3%, go back to step 1.

12. If mole fraction CO in dry gas is 3%, mass of catalyst is found as number of

increments multiplied by 200lbs.

The MATLAB program can be found in Appendix 4.12.1.



Calculations

The temperature of inlet stream to the HTS reactor was varied to get the corresponding mass

of catalyst required to reduce CO to 3% (dry basis). This was done to optimize the mass of

catalyst used with its optimum inlet temperature of the inlet stream. Then, a graph was

plotted to show the relationship between mass of catalyst and inlet temperature.

Fig. 4.7.1.1 Graph of mass of catalyst against inlet temperature

From the graph, the minimum mass of catalyst needed was 121909 kg.

However, 5% more catalyst was added to allow for any degrading of catalyst.

Therefore mass of catalyst used = 1.05×121909kg = 128000kg

This occurred at the inlet temperature of 627 K. Thus, for the HTS reactor, the inlet stream

was fixed at 627 K.



A temperature profile graph was also plotted in the figure below. The temperature range was

from the inlet stream of 627 K to that of the outlet stream was which calculated to be 693.2 K.

Fig. 4.7.1.2 Temperature profile graph



A conversion profile graph was also plotted as shown below. The inlet CO was at 13.34% dry

basis and the outlet CO was 3% dry basis.

Fig. 4.7.1.3 Conversion profile graph

4.7.2 Pressure drop

Pressure Drop, though negligible in some reactors, can be a major concern in others. It is an

important variable in the rate equations for gaseous reactions. Since compressors and

compressor operating costs often dominate the economic structure of a reactor system,

pressure drop is not only important but must be predicted with good accuracy.

The resulting force must not exceed the crushing strength of the particles. In homogeneous

clean beds, one would expect the maximum stress to occur at the bottom of the bed, where

the weight of the catalyst combines with the stress created by the ∆P across the bed. In down

flow, this force created by the ∆P is transmitted by the contacting solids to the bottom of the

bed. Some catalysts are quite fragile and this issue demands close attention with sufficient

safety factor applied.



Mass velocities through the bed must be high enough to minimize inter-phase gradients and

assure good distribution. Incremental increases in ∆P, however, should not cause pumping or

compressing costs to exceed savings realized from improved reactor performance. In many

packed bed systems, the maximum economical ∆P is in the range of 3-15 % of the total

pressure.

In gaseous systems, higher pressure drop and thus higher velocity, also means smaller

diameter reactors, which can be important in reducing costs of high pressure reactors; but this

advantage can be offset by higher energy costs. The given fraction of plant pressure drop

allotted to drop across the bed is directly proportional to the fraction of power consumed,

which is essentially a function of energy costs and independent of total pressure. Thus

economic allowable ∆P will be a fixed fraction of total pressure and can vary from a few

inches of water for reactors operating near atmospheric pressure to several atmospheres for

reactors operating at higher pressure.

A unique value of particle density does not even exist for a given catalyst. Generally, smaller

sizes will have higher particle densities than larger sizes, which can be rationalized by

considering the limit of a catalyst approaching the size of an average pore. Dense packing in

a full-size bed is preferred for uniform flow distribution and is obtained by raking or

spreading the catalyst between each load. Although a rapidly dumped bed will result in looser

arrangement and lower pressure drop, it is more likely to cause channelling.

Although small catalyst particles have higher effectiveness factors, it is not wise to specify

sizes below 1/8 in. unless some means is provided for removing fines, dirt and scale from the

feed stream. The greatest care should be exercised in packing a bed to eliminate fines and dirt

and the reactor should be protected by suitable filters whenever plugging by scale or polymer

formation in upstream equipment is anticipated. These materials can be carried to the reactor

and deposited on the top part of the bed and limit the throughput drastically. Plugging of a

catalyst bed is a serious problem that can ultimately lead to shutdown and dumping of the bed

as pressure drop becomes excessive. Prior to this event, serious malfunction of the reacting

fluid can occur, resulting in poor yields and reduced production. The ability to predict clean-

bed ∆P is often foreshadowed by our inability to predict the rate and character of plugging

that may occur.



Methodology

For this HTS reactor, the inlet pressure is 2610kPa. Since maximum economical pressure

drop is in the range of 3-15% of the total pressure. [3] Taking pressure drop to be 4% of total

pressure.

∆P = 0.04×2610 kPa

= 104.4 kPa

≈ 1 bar

Therefore, a pressure drop of 1 bar is to be obtained.

These are the equations required to find the pressure drop and aspect ratio.

µ

GDN

p=Re ,

Where NRe = Reynolds’ number

c

c

c

p

h

d

dD

24

6

+= = 0.25

µ = Average of inlet and outlet viscosity

Re

115075.1

Nf k

ε−+= ,

Where fk = friction factor

ε = voidage

LgD

GfP

cfp

k

−=∆

3

21

εε

ρ,

Where G = mass flux

= mass flowrate per cross sectional area

fρ = density of feed

L = length of reactor



To find the aspect ratio of the reactor which corresponds to a pressure drop of 1 bar, the

following algorithm was used.

1. The volume of the total catalyst was calculated based on mass of catalyst used and

density. This volume is multiplied by 1.2 to give an extra 20% volume for the

allowance of inert support as well as poor packing of catalyst.

2. Calculate average µ (viscosity) based on inlet and outlet.

3. Assume a value of L

4. Calculate the corresponding value of diameter,

=L

VD

π2

5. Calculate aspect ratio, AR=L/D

6. Calculate G

7. Calculate ε

8. Calculate N Re

9. Calculate fk

10. Calculate ∆P

11. Go back to step 3 and assume another value of L to get corresponding ∆P .

12. Plot graph of ∆P against AR.

13. Identify the AR where ∆P =1 bar

The MATLAB program can be found in Appendix 4.12.2.



Calculations

The pressure drop against aspect ratio figure is plotted as shown below:

Fig. 4.7.2.1 Graph of pressure drop against Aspect ratio

Therefore, a pressure drop of 1 bar corresponds to an Aspect ratio of 3.497.

Solving simultaneous equations 497.3=D

L-------------- (1)

V= Mass of catalyst/ bulk density

V= 113.92m3

92.1132

2

=

L

Dπ --------(2)

L, length of reactor = 12.11m

D, diameter of reactor = 3.46m



4.7.3 Thickness of vessel

With an internal pressure of 2610 kPa and diameter of 3.46m=11.35ft,

Design Pressure: 2610 x 1.2 x 0.145 = 454.2 psi, taking into account an allowance of 20% for

increased operating pressure

Design Temperature: 890 0F (maximum catalyst use temperature)

S = 13100 psi

E, the joint efficiency = 1.0, for double butt welded and fully radio-graphed welds

Minimum Corrosion allowance = 1/8 in.

413.2)2.4546.0()0.113100(

1268.52.454

6.0=

×−×

××=

−=

PSE

PRt in

538.2125.0413.2 =+=actualt in

The MATLAB program used to solve this can be found in Appendix 4.12.3.

4.7.4 Reactor size and cost

Catalyst cost

Mass of catalyst = 128000 kg

Cost of catalyst = US$20/cu ft in 1977

Particle density of catalyst = 126 lb/cu ft

Volume of catalyst = 2234.9 cu ft

CEPCI in 1977 = 204.1

CEPCI in 2006 = 499.6

Cost of catalytic bed = 2234.9 x $20 x 499.6/204.1

= US$109,414



Vessel Costs

From the literature data [1], the vessel cost can be estimated as follows:

Using 2:1 elliptical heads of same thickness

Material Density: ( 490=ρ lb/ cu ft)

Shell, =×××× 49072.39)12

538.2(35.11π 146,778 lb

Heads

=××+××× 212

538.2]

12

538.2)35.1123.1[(

4490 2π

32,696 lb

Total Weight = 146,778 + 32,696 = 179,474 lb

For this size and type vessel, a fabricated cost of 73 cents/lb without nozzles was suggested

as an estimating figure (1971 cost) by a fabricator.

CEPCI in 1971 = 132.3

CEPCI in 2006 = 499.6

Cost of vessel= 750,494$3.132

6.49973.0179474 US=××



4.8 HEAT EXCHANGER

With a conversion of 75.27% CO, a conversion reactor is simulated in Hysys. This is the

figure that is extracted from Hysys to get the duty needed to cool down the HTS outlet before

it goes into the LTS inlet. The LTS inlet is to be cooled to 493.1K.

Fig. 4.8.1 Hysys diagram of HTS reactor

The duty needed to cool the HTS outlet stream from 692 K to 493.1K is 1.099e+008 kJ/hr

according to Hysys. However, the actual HTS outlet temperature as calculated from

MATLAB is 693.2K. This is quite close to the calculated value from Hysys. The actual duty

would be further discussed in the next part, Chapter 4.8.1.



4.8.1 Heat Exchanger Design Considerations

Shell and tube heat exchanger which is the most commonly used basic heat exchanger

configuration in the process industries is selected because it provides a comparatively large

ratio of heat transfer area to volume and weight and it is mechanically rugged enough to

withstand normal shop fabrication stresses and normal operating conditions. Also, it can be

easily cleaned and components susceptible to failure (gaskets and tubes) can be easily

replaced.

According to the heat exchanger network design, there are 3 heat exchangers designed to cool

down the HTS outlet stream to the cooled LTS feed stream. I will be designing the heat

exchanger which is used to cool HTS outlet and to heat up SMR feed.

Split ring internal floating head heat exchanger is selected for this heat exchanger. It can be

used for liquids that foul as the tubes and bundle can be removed from shell for cleaning or

repairing without removing the floating head cover. Since the HTS outlet is the stream that

causes more fouling than the SMR feed, I have chosen to use HTS outlet in the tube side, and

SMR feed in the shell side.

4.8.1.1 Physical properties extraction

The physical properties of the two streams are extracted from Hysys. The average values are

used for the design. Where duty, Q = 4.622 MW.

HTS outlet inlet outlet mean

temperature (o C) 263.3 234.3 248.8

specific heat(kJ/kg-C) 2.637 2.622 2.629

thermal conductivity (W/m-C) 0.1123 0.1072 0.1097

density(kg/m3) 7.444 7.893 7.661

viscosity(cp) 0.01759 0.01700 0.01735



SMR feed

temperature (o C) 218 253.3 235.7

specific heat(kj/kg-C) 2.326 2.342 2.334

thermal conductivity (W/m-C) 0.04511 0.0491 0.04712

density(kg/m3) 12.60 11.57 12.09

viscosity(cp) 0.01424 0.01535 0.01480

4.8.1.2 Determination of overall heat transfer coefficient

After iterations, the overall heat transfer coefficient was U=540.1 W/m2oC. For an exchanger

of this type with light gases as hot gas and methane and water vapor as cold gas, the overall

heat transfer coefficient according to Table 12.1 of Coulson and Robertson’s Chemical

Engineering Design textbook falls in the acceptable region.

4.8.1.3 Exchanger type and dimensions

=∆ mT Shell can be carbon steel. Tube can be stainless steel due to H2 pitting. The HTS outlet

is dirtier than the SMR feed, therefore put the HTS outlet through the tubes and the SMR feed

through the shell.

C

TT

TT

TTTTT

incouth

outcinh

incouthoutcinh

LMTD

Ο=

−

−−−−

=

−

−

−−−=∆ 89.12

2183.234

3.2533.263ln

)2183.234()3.2533.263(

ln

)()(

,,

,,

,,,,

8215.02183.253

3.2343.263=

−

−=R , and 7792.0

2183.263

2183.253=

−

−=S

From Fig 12.19, 70.0=tF , which is acceptable.

CTm

Ο=×=∆ 02.989.1270.0



4.8.1.4 Heat transfer area

23.91502.91.540

6^104622.4m

TU

QA

m

=××

=∆×

=Ο

4.8.1.5 Layout and tube size

A split-ring floating head exchanger is used for efficiency and ease of cleaning. Use

19.00mm outside diameter, 15.00mm inside diameter, 5m long tubes on a triangular

23.80mm pitch.

4.8.1.6 Number of tubes

Area of one tube(neglecting tube sheets thickness) = 51000.19 3 ××× −π = 0.2985m2

Number of tubes = 915.3 / 0.2985 = 3066

So, for 2 passes, tubes per pass = 3066 / 2 = 1533

(Check for tube-side velocity to see if reasonable)

Tube cross-sectional area = (π /4)(15 ×10-3)2 = 0.0001767 m2

Thus, area per pass = 1533× 0.0001767 = 0.2709m2

Volumetric flow = (2.074×10^5/3600) × (1/7.661)=7.52 m3/s

Tube side velocity, tu = 7.52 / 0.2709=27.76 m/s

4.8.1.7 Bundle and shell diameter

For 2 tube passes, K1= 0.249, n1= 2.207,

So, Db = 19.0× ( 1533 / 0.249 )1/ 2.207 = 1.36 m

For a split-ring floating head exchanger the typical clearance is 20 mm, so the inside shell

diameter, Ds= 1.36 + 0.02 = 1.38 m



4.8.1.8 Tube-side heat transfer coefficient

Re = =×

×××−

−

3

3

1001735.0

101576.27661.7 183,801

Pr = 1097.0

1001735.010629.2 33 −×××= 0.42

=iD

L

00.15

5000 = 333

From figure 12.23, jh = 0.045

Nu = 281)42.0()183801(023.0 33.08.0 =×

hi = 20491000.15

1097.0281

3=

××

− W/m2C

4.8.1.9 Shell-side heat transfer coefficient

Take baffle spacing to be Ds /5 = 1.38/5 = 0.276 m = 276mm. This spacing should give good

heat transfer.

As= 223 151.01015113804.0138080.23

00.1980.23mmm =×=××

−

De= ( ) mm76.1800.19785.080.2300.19

27.1 22 =×−

Volumetric flow rate = 1.835×105/3600/12.085 = 4.217m3/s

Shell-side velocity, us= 4.217/0.151 = 27.9 m/s

Re = =×

×××

−5

5

1048.1

0188.0151.03600

10835.1

4.274×105

Pr = 04712.0

1001480.010334.2 33 −×××= 0.733

From Fig 12.29, jh=0.45

533.05 1036.4733.010274.445.0)01876.0

04712.0( ×=××××=sh W/m2C



4.8.1.10 Overall coefficient

436000

1

6.5678

1

552

15

19ln1019

15

19

1249

1

2049

11

3

0

++×

××+

+=

−

U

0U =540.1W/m2 oC.

4.8.1.11 Pressure drop

kPaPaPt 722722739)751.27661.75.0()5.2015.0

5045.08(2 2 ==×××+×××=∆

From Fig 12.30, jf = 0.028

kPaPaPs 1851854182

9.27085.12

5.0

83.4

0188.0

37.1028.08

2

==×

××××=∆



4.9 CONCLUSION

The HTS reactor designed has determined the parameters as shown:

1. Conversion of CO from 13.34% dry basis to a composition of 3% dry basis

2. The mole fraction of the outlet of HTS

3. The weight of catalyst needed for the reaction.

4. Pressure drop of the reactor.

5. The dimensions of the reactor were also calculated, namely length of reactor, diameter

of reactor and thickness of reactor.

6. The cost of the catalyst needed was calculated, as well as the cost of the vessel.

These give a good idea on the design of the HTS reactor as well as the cost of building the

reactor.



4.10 NOTATIONS

iν : stoichiometric coefficient of species i

iµ : chemical potential of species i

n : number of moles

G : specific gibbs energy

ε : reaction coordinate

T : temperature of system

P : pressure of system

tG : total gibbs free energy

∏i

: product over all species i

ν : total stoichiometric number

)( COr− : reaction rate in lb moles CO converted/(lb catalyst/hr)

ψ : activity factor

k : rate constant

K : equilibrium constant

iy : mole fraction of species i

bρ : bulk density of catalyst (lb/ft3)

T : temperature in K

F : component molar flow rate

W : weight of catalyst

cp,j : heat capacity of component j

t : minimum thickness of wall without corrosion

P : design pressure of the reactor vessel

R : internal radius of shell without corrosion

S : maximum allowable stress value

E : joint efficiency (assume = 1)

Dp : characteristic length of pellet (ft)

G : mass flux (lb/s)

µ : average viscosity of fluid (cP)



NRe : reynolds number

L : length of reactor

D : diameter of reactor

dc : diameter of cylindrical catalyst pellet (ft)

hc : height of cylindrical catalyst pellet (ft)

ρb : bulk density of catalyst pellet (70 lb/ft3)

ε : voidage

fk : friction factor

ρf : density of fluid (lb/ft3)



4.11 REFERENCES

1. H.F. Rase, Chemical Reactor Design for Process Plants, Vol 2, New York Wiley,

1977

2. J.M. Smith, H.C. Van Ness, M.M. Abbott, Introduction to Chemical Engineering

Thermodynamics 7th ed, 2005

3. H.F. Rase, Chemical Reactor Design for Process Plants, Vol 1, New York Wiley,

1977

4. R. Q. Barr, A Review of Factors Affecting the Section of Steels for Refining and

Petrochemical Applications, Climax Molydenum Co., Greenwich, Conn., 1971

5. C.H. Samans, Hydrocarbon Process., 42(10), 169 and (11) 241, 1963

6. H. M. Spencer, Industrial Engineering Chemistry, Vol 40, pg 2152-2154, 1948

7. K. K. Kelley, U.S Bur.Mines Bull. 584, 1960

8. L. B. Pankratz, U.S. Mines Bull. 672, 1982



4.12 APPENDICES

Appendix 4.12.1

T= input('Please enter initial temperature for reactor in K:'); F= 2.2*15797; %flowrate in lbmol/hr yco= 0.080035; yh2o= 0.400173; yco2= 0.049934; yh2= 0.43523; yc2h6= 0.001975; ych4= 0.031995; yn2= 0.000658;

Fco= F*yco; Fh2o= F*yh2o; Fco2= F*yco2; Fh2= F*yh2; Fc2h6= F*yc2h6; Fch4= F*ych4; Fn2=F*yn2;

k=exp(15.95-4900/T); K=exp(-4.33+4578/T); rate=4*k*(yco*yh2o-yco2*yh2/K)/(379*70);

Fco1= Fco -rate*200; Fh2o1= Fh2o-rate*200; Fco21= Fco2 + rate*200; Fh21= Fh2 +rate*200;

cpco= (3.376+(0.557/10^3)*T -(0.031*10^5)*T^(-2))*8.314; cpc2h6=(1.131+(19.225/10^3)*T-(5.561/10^6)*T^2)*8.314; cph2o=(3.47+(1.45/10^3)*T +(0.121*10^5)*T^(-2))*8.314; cpco2=(5.457+(1.045/10^3)*T -(1.157*10^5)*T^(-2))*8.314; cph2=(3.249+(0.422/10^3)*T +(0.083*10^5)*T^(-2))*8.314; cpn2=(3.28+(0.593/10^3)*T +(0.04*10^5)*T^(-2))*8.314; cpch4=(1.702+(9.081/10^3)*T -(2.164/10^6)*T^2)*8.314;

deltaa = 5.457 + 3.249 - 3.376 - 3.470; deltab = (1.045 + 0.422 - 0.457 - 1.450) * 10^-3; deltad = (-1.157 + 0.083 - (-0.031) - 0.121) * 10^5;

integral =(deltaa * 298.15 * ((T/298.15) - 1) +

(deltab/2)*(298.15^2)*(((T/298.15)^2)-1) + (deltad/298.15)*(((T/298.15) -

1)/(T/298.15)));

dHco = -41166 + 8.314 * integral;

dT = ((rate * 200) * (-

dHco))/(Fco*cpco+Fh2o*cph2o+Fco2*cpco2+Fh2*cph2+Fc2h6*cpc2h6+Fn2*cpn2+Fch4*

cpch4);



T1 = T + dT;

yco1 = Fco1/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6); yh2o1= Fh2o1/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6); yco21= Fco21/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6); yh21= Fh21/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

ydryco= (yco1/(1-yh2o1));

i=1;

while ydryco>0.03,

k=exp(15.95-4900/T); K=exp(-4.33+4578/T); rate=4*k*(yco*yh2o-yco2*yh2/K)/(379*70);

k=exp(15.95-4900/T1); K=exp(-4.33+4578/T1); rate1=4*k*(yco1*yh2o1-yco21*yh21/K)/(379*70);

rateavg= (rate+rate1)/2;

Fco = Fco1; Fh2o = Fh2o1; Fco2 = Fco21; Fh2= Fh21;

Fco1= Fco -rateavg*200; Fh2o1= Fh2o-rateavg*200; Fco21= Fco2 + rateavg*200; Fh21= Fh2 +rateavg*200;

cpco= (3.376+(0.557/10^3)*T1 -(0.031*10^5)*T1^(-2))*8.314; cpc2h6=(1.131+(19.225/10^3)*T1-(5.561/10^6)*T1^2)*8.314; cph2o=(3.47+(1.45/10^3)*T1 +(0.121*10^5)*T1^(-2))*8.314; cpco2=(5.457+(1.045/10^3)*T1 -(1.157*10^5)*T1^(-2))*8.314; cph2=(3.249+(0.422/10^3)*T1 +(0.083*10^5)*T1^(-2))*8.314; cpn2=(3.28+(0.593/10^3)*T1 +(0.04*10^5)*T1^(-2))*8.314; cpch4=(1.702+(9.081/10^3)*T1 -(2.164/10^6)*T1^2)*8.314;

integral = (deltaa * 298.15 * ((T1/298.15) - 1) +

deltab/2*298.15^2*((T1/298.15)^2-1) + deltad/298.15*(((T1/298.15) -

1)/(T1/298.15)));

dT = ((rateavg * 200) * (-

dHco))/(Fco*cpco+Fh2o*cph2o+Fco2*cpco2+Fh2*cph2+Fc2h6*cpc2h6+Fn2*cpn2+Fch4*

cpch4);



yco = yco1; yh2o = yh2o1; yco2 = yco21; yh2 = yh21;

yco1 = Fco1/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6); yh2o1= Fh2o1/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6); yco21= Fco21/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6); yh21= Fh21/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

ydryco= yco1/(1-yh2o1);

x(i) = i*200/2.2; y(i) = T; a(i) = ydryco;

T = T1; T1 = T + dT; i= i + 1; end

Wt=i*200/2.2;

fprintf('mass of catalyst is %f kg.\n',Wt); fprintf('outlet temperature of HTS is %f K.\n',T1); fprintf('mol fraction of CO is %f .\n',yco1); fprintf('mol fraction of H2O is %f .\n',yh2o1); fprintf('mol fraction of CO2 is %f .\n',yco21); fprintf('mol fraction of H2 is %f .\n',yh21); fprintf('mol fraction of N2 is %f .\n',yn2); fprintf('mol fraction of CH4 is %f .\n',ych4); fprintf('mol fraction of C2H6 is %f .\n',yc2h6); fprintf('mass flowrate in lb/hr is %f .\n',F);

plot (x,y); plot (x,a);

Appendix 4.12.2

mass = 281609.79; %mass of catalyst in lbs flowrate = 456300; %flowrate in lbs per hour u1 = 0.01999; %inlet viscosity in cp u2 = 0.02118; %outlet viscosity in cp pf = 0.4125; %density of feed in lbs per feet3 u = (u1+u2)/2; V = 1.2*mass/70;

for i=1:600; L = i*0.1; D = 2*((V/(pi*L))^0.5);

AR = L/D; area = pi*(D/2)^2; G = flowrate/area;



Dp = 0.25/12; Nre = Dp*G/(2.42*u); fk = 1.75+150*(0.555)/Nre; dPft = L*((fk*G^2/(Dp*pf*32.17*(3600^2)))*0.555/(0.445^3)); dPsi = dPft/(12^2); dPbar = dPsi/14.7;

x(i,1) = AR; x(i,2) = dPft; x(i,3) = dPsi; x(i,4) = dPbar; x(i,5) = G;

end plot(x(:,1),x(:,4));

Appendix 4.12.3

mass = 281609.79; % mass of catalyst in lbs AR= 3.497; %AR P = 454.2; %pressure in psi S = 13.1; %maximum allowable stress value in kips per inch square

V = mass*144*12/70; E = 1; D = (4*V/(pi*AR))^(1/3); L = D*AR;

t=P*(D/2)/(S*1000*E-0.6*P); D=D*0.0254; L=L*0.0254; fprintf ('t= %f inches \n',t); fprintf ('D= %f metre \n',D); fprintf ('L= %f metre \n',L);

CN 4120: Design II Team 32: Tham Zhi Yong, Andrew (U046754W) LTS Unit Design Report


Chapter 5 : LOW TEMPERATURE SHIFT REACTOR

5.1 INTRODUCTION

The water shift reaction usually occurs in an fixed bed adiabatic system with the presence of

a catalyst to speed up the reaction rate. In an adiabatic system, CO slip is determined by the

exit temperature of the shift converters, because low temperatures results in low equilibrium

levels of CO, as the following exothermic process is taking place:

CO + H2O CO2 + H2 ∆∆∆∆H = -41.2kJ/mol (5-1)

On the other hand, favorable kinetics occurs at higher temperatures. Either a high steam-to-

gas ratio or low temperature can be used to improve CO conversion percentage, but that also

correspondingly contribute to higher capital and operation cost. Hence there is a tradeoff

between CO conversion percentage and costs.

Fig.1 Typical CO variation in high temperature and low temperature shift catalyst beds

[Frank, 2003a]

Conversion in a single high-temperature-shift(HTS) converter is equilibrium limited. Since

this reaction is exothermic, the rise in temperature as reaction proceeds will eventually not

favor further reaction. This limitation can be overcome by employing a second converter, the

low-temperature-shift (LTS) converter after the HTS converter. Usually an inter-bed cooling

process is employed between the two converters to keep the reaction occurring at low

temperature in the second converter. A knock-out drum is then employed to condense and

remove all water prior to feeding into the pressure swing adsorption (PSA). This part of the

design project presents detailed chemical engineering design of a LTS converter and the

knock-out drum.



Fig.2 LTS converter in the HYSYS environment

5.2 LTS DESIGN CONSIDERATIONS

5.2.1 Current Status

Carbon monoxide exits the HTS converter with a molar fraction of 0.03(dry basis) at 420oC.

The stream, with a molar flow of 11354kmol/h, is then cooled to bring the temperature down

to 220oC before feeding into the LTS converter. Molar composition of the feed is illustrated

in the following diagram.

Fig.3 Molar compositions of feed into LTS converter

The outlet composition was automatically generated using Hysys, using the rate equation

associated with this reaction (and catalyst type). However, it should be noted due to its

iterative nature, Hysys could not obtain a value closer to that of the exact situation than

Matlab. Hence, there is still a need to carry out interations (based on the same rate equation)

in the Matlab environment. A comparison between results calculated from both programs will

be made in latter section.



5.2.2 Kinetics of Low-Temperature Water-Gas-Shift (LTWGS)

Equation (5-1) may be represented in the following form

A + B C + D (5-2)

where A, B,C and D are CO, H2O, CO2 and H2 respectively

Rase (1977) has come up with the following equation for application to the shift conversion:

(5-3)

Where Xi = the dimensionless concentration of component i (Ci/Cref)

k = rate constant

= exp (12.88 -1855.6/T) for copper-zinc catalyst

K = equilibrium constant

= exp (-4.72 + 8640/T) for 760 ≤ T ≤ 1060

P = pressure, atm

(-rco) = rate, lb moles CO converted/(lb catalyst)(hr)

T = temperature, K

yj = mole fraction of component indicated

ρb = catalyst bulk density, lb/cu ft

ψ = activity factor for the copper-zinc catalyst

Copper-zinc catalyst ψ = 0.86 + 0.14P for P ≤ 24.8

= 4.33 for P ≥ 24.8



5.2.2.1 Assumption made for equation (5-3) :

This equation represents the activity level characteristic of mid-life of the catalyst.

These rate constants have been expressed on the basis of a reasonable “lined-out” activity that

the catalyst would maintain for a considerable time provided operating errors which cause

deactivation do not occur.

Multiplying the rate equation by ρb , we obtain the rate of reaction in units of moles of CO

converted per unit volume of catalyst per second, and converting the units to S.I units, we

obtain the following equation

(5-4)

Where

(5-5)

The pre-exponential factor ko includes the diffusion effect as given by the catalyst

manufacturer (Rase, 1977).



5.2.3 LTS Catalyst

From (5.3), it is important to decide the catalyst which we are using for the LTS shift.

Copper-Zinc Oxide supported on alumina will be taken as the catalyst for our design and its

specifications will be used for the calculations.

Copper-Zinc oxide offers the thermodynamic advantage of a lower operating temperature for

the exothermic reaction in eq. (5-1)

Characteristics of the catalyst assumed for this design are as follows:

Catalyst Type Copper-Zinc Oxide supported on alumina

Maximum Operating Temperature (oC) 260 - 288

Tablet Size (in.) ¼ x 1/8

Bulk Density (lb/cu ft) 90

Particle Density ($/cu ft) 155

Cost ($/cu ft) 75

Catalyst Poison Sulfur and halogen compounds, as well as

unsaturated carbons

Catalyst life 2-3 yr

Fig.4



5.2.3.1 Characteristics of the industrial LTS catalyst

The low-temperature shift catalyst is usually a mixture of copper oxide and zinc oxide in a

ratio between 1:1 1:2, with alumina added in place of some of the zinc oxide. In addition,

promoters such as Cr2O3, MnO, or some metal oxides have been used. Chromium oxide has

been used in place of alumina. Preparative procedures on the whole are more critical for the

LTS catalyst than for the HTS catalyst.

5.2.3.2 Preparation

Preparative procedures are much more critical for the LTS catalyst as compared to the HTS

catalyst. Coprecipitation of the metals as metal nitrates are carried out via pH adjustment with

ammonium bicarbonate. The oxides formed in this way are intimately intermixed by this

procedure, which is essential for high activity and stability. It has been suggested that ZnO in

excess can protect the copper content from inadvertent sulfur poisoning. Aluminum oxide

also serves as a stabilizer for the copper, preventing it from being sintered easily. Thus we

can see that the manufacture of these LTS catalysts involve great skill and refined proprietary

techniques.

Since the LTS catalyst is pyrophoric, it must be sequested during system shutdown when

only air flows through the system.

5.2.3.3 Supply

Catalyst suppliers usually offer thorough instructions for the start-up, catalyst reduction,

operation and shutdown for the particular catalysts purchased. Instructions for catalyst

reduction and start-up are particularly critical, since excessive temperatures must be avoided.



One of the many suppliers for the LTS catalyst is Haldor Topsoe. Listed below are

properties of an example of the LTS catalyst produced from this company.

http://www.topsoe.com/

5.2.3.4 Deactivation of LTS Catalyst

(I) Poisons

Common poison of the LTS catalyst are sulfur and chlorine compounds. Sulfur compounds

such as H2S are removed in the ZnO adsorber beds prior to feeding into the steam-methane

reformer. However there could exist times of upset such as short-periods of high-sulfur feed.

In such instance, break-through sulfur will occur and pass to the HTS converter. There is a

high possibility the HTS catalyst will be able to safely adsorb the H2S and protect the LTS

bed. In some circumstances, sulfur may still get into the LTS. It is for this reason that LTS

catalysts contain excess ZnO so that upper portion of the bed can serve as a sulfur guard. Zinc

sulfide forms for the early part of the bed but further down the bed, sulfur in the form of H2S

is chemisorbed, the extent of this happening depends on the operating operations.

Consequently, deactivation of actives sites will take place.

Chloride compounds are a major and permanent poison of LTS catalyst, and worth a mention

despite the fact that no chloride compounds are involved in this design project.



CuCl and ZnCl2 are formed and they can cause copper crystal growth (sintering) and

significant loss of catalyst activity. Ways to tackle the problem of choride poisoning involve

installing a bed of chlorine adsorbent (e.g. CaO/ZnO or alkalized alumina) upstream of the

ZnO adsorbent bed prior to the reformer, another bed is placed above the LTS catalyst,

composed of a chlorine adsorbent as well.

(II) Sintering

Excessive temperature can result in sintering. Very small crystallites of copper are

thermodynamically favored to coalesce into large crystals and thus produce a less active

catalyst due to low porosity overall. Despite the fact that these crystallites are stabilized by

the associated ZnO and also alumina, this protection is destroyed at elevated temperatures.

Inlet operating temperatures for LTS between 175-275oC have been suggested, but it is

always encouraged to operate at the lowest temperature possible, since sintering is a

phenomenon related to both time and temperature. However, there is a lower temperature

limit for the operating condition in the LTS converter, to avoid the any condensation of the

steam we use in the low temperature shift reaction, as any condensation in the pores can

result in catalyst damage. It is often suggested that the lowest temperature should be no lower

than 20oC above the dewpoint.

5.2.3.5 LTS catalyst in operation

A common practice of some hydrogen-producing companies is to increase temperature

during a LTS operating cycle to overcome deactivation of the catalyst. This will however,

increase the growth of the crystals and shorten the life of the catalyst. In the lower regions of

temperature, sintering rate is very low but this increases as temperature is raised. Ultimately,

deactivation rate becomes significant and the catalyst activity will suffer. The higher the bed

temperature reached, the more critical temperature control becomes, particularly if the

process gas is introduced after the bed reaches operating temperature. Sudden rapid rise in

temperature can damage the catalyst.

Reducing gas of H2 mixed with N2 or with natural gas is usually recommended for LTS

catalyst to keep the catalyst in the reduced form.



5.2.3.6 Assumptions made for LTS Catalyst

The following assumptions have been made for the derivations of the differential equations

which characterize the pellet mass and heat balance

1. The Copper Zinc catalyst pellet particles have a homogenous porous structure

2. Mass transfer within the catalyst particles occurs by diffusion only which may be

expressed by a constant effective diffusion coefficient D, and rate of intraparticle

diffusion is described by Fick’s Law

3. Conduction is responsible for the thermal transfer within catalyst particles and

effective thermal conductivity λe is used with the Fourier’s law, to describe the

intraparticle heat conduction

4. Both mass and heat transfer within the catalyst pellets only take place in the radial

direction

5.2.3.7 Mass balance on the Copper-Zinc catalyst pellet

Dimensionless steady state material balance for component I over a shell of dimensionless

thickness dw ( where w is the dimensionless radial coordinate, z/Rp) is given by:

(5-6)

where

(5-7)

φi is Thiele’s modulus of the pellet, for component i and is defined as

(5-8)

and

(5-9)



γ is the dimensionless activation energy, defined as

(5-10)

Equation (5-10) is a second order differential equation of the boundary value type having two

split boundary conditions:

At ω = 0

(5-11)

At ω = 1

(5-12)

where i = A,B,C,D

5.2.3.8 Heat balance on the Copper-Zinc catalyst pellet

The dimensionless steady state enthalpy balance over the shell of dimensionless thickness

dω is given by

(5-13)

Where βi is the thermicity factor of the pellet based on component I, defined as

(5-14)

The boundary conditions being:

At ω = 0

(5-15)

At ω = 1

(5-16)



The non-isothermal effectiveness factor for a spherical particle the non-isothermal

effectiveness factor η is defined as:

(5-17)

In dimensionless form the above equation becomes

(5-18)

It can also be written as

(5-19)



5.2.4 Modeling the converter

5.2.4.1 Assumptions made for the converter

The heterogeneous model is developed in terms of bulk variables with the effectiveness

factor introduced to account for the diffusional limitations. The assumptions made for the

overall reactor model are as follows:

1. There is uniform distribution of gas flow velocity inside the converter

2. The reactor is studied under steady state conditions

3. The radial distribution of the temperature and concentration of the different

components inside the converter is uniform, i.e. the model is one-dimensional

4. Heating and mass diffusion in the longitudinal direction are negligible considering the

very high gas velocities at which the reactor is operated, i.e. axial dispersion is

negligible

5. The pressure drop across the reactor is negligible compared with the total pressure of

the reactor

5.2.4.2 Reactor mass balance

For the bulk gas phase, the rate of reaction is formulated in terms of the mole fractions Yi

instead of the dimensionless concentrations Xi (as in the catalyst pellets equations). This is a

more convenient approach as the total number of moles is constant, while the volumetric gas

flowrate is changing due to the change of temperature.

At steady state, a component mass balance on CO over an element of catalyst bed of

thickness dl and a cross sectional area Ai , with a constant total molar flow rate nT , gives

(5-20)

where nA is the molar flow rate of component A, and the rate of reaction is given by (Rase,

1977)

(5-21)



In a dimensionless form, the equation becomes

(5-22)

Where z’ = I/L and

KB is the temperature dependent equilibrium constant which is defined as (Borgars and

Campbell, 1974):

(5-23)

where yB = TB/Tref and (CO2), (H2), (CO), (H2O) are the partial pressures or fugacities for the

different species in equilibrium.

The boundary condition at the inlet of the reactor is at z’=0

(5-24)

5.2.4.3 Reactor mass balance

At steady state, the heat balance equation is

(5-25)

In a dimensionless form, the equation will be

(5-26)

where

(5-27)

The boundary condition is at z’= 0

(5-28)

The bulk phase temperature can also be computed from the bulk phase concentration by

making a cumulative heat balance over any reactor length

(5-29)



Therefore, it is sufficient to integrate (5-22). There is no need to integrate equation (5-27)

along the length of the reactor, equation (5-29) can be used instead.

From stoichiometry, at any depth of the reactor, the concentration of CO2, H2O, H2 can be

expressed in terms of the bulk concentration as follows:

(5-30)

Where I = B, C, D and a =-1 for reactants and a = +1 for products.

5.2.4.4 Transport parameters

Viscosity and thermal conductivity

The viscosity µ of the fraction, designated by subscript x at a density ρ and temperature T is

given in terms of a reference fluid, designated by subscript o. The equation is

(5-31)

where

(5-32)

With M as the molecular weight and To, ρο defined by the ratios

(5-33)

and

(5-34)

fx,o and hx,o are scaling ratios, which are in general

(5-35)

and

(5-36)



The subscript c denotes the critical value and the superscript * denotes reduction of the

variables by the critical value. The functions θ and φ are the shape factors expressed in

terms of Pitzer acentric factor, ω , via functions of the form

(5-37)

and

(5-38)

where F and G are universal functions reported for example, by Leach et al. (1968) and Ely

and Hanley (1981). The thermal conductivity λ is also evaluated through the same procedure

(Ely and Hanley, 1981)

Prandtl number

The prandtl unmber Pr is computed as

(5-39)

Diffusion coefficients

The binary diffusion coefficient of each component is computed by the relation

(5-40)

where vi are the values of the atomic and structural diffusion-volume coefficients (Perry et

al,1984)

The value of the diffusion of each component in the mixture is calculated by the relation

(Bird et al. 1960)

(5-41)

where Yi = mole fraction of each component, Dij = the binary diffusion coefficient, and Di,mix

= diffusion coefficient for each component in the mixture.



External mass and heat transfer coefficients

Correlations for both mass and heat transfer coefficients kg and h are found empirically from

the mass and heat transfer J-factor (JD and JH) correlations, which are defined as

(5-42)

and

(5-43)

The values for JD and JH are almost equal and are computed as a function of the Reynolds

number:

JD = JH = 0.989 Re-0.41 for Re>350

= 1.820 Re-0.51 for Re<350 (5-44)

where, Re = Reynolds number = Gdp / µ µ µ µ

The external mass transfer in the pellet equation takes the form

(5-45)

and the external mass transfer coefficient is calculated by the relation

(5-46)

where RG= the universal gas constant and T = average temperature for the catalyst bed.



5.3 DESIGN CONDITIONS

5.3.1 Temperature

Inlet temperature has been set at 220oC, a temperature which is low enough to prevent the

sintering of the LTS catalyst, and high enough to prevent any condensation of steam taking

place, damaging the catalyst as a result. Referring to (5-3), as pointed out by Rase, it can

shown that rate doubles for a rise of 200oC for the LTS catalyst. By using around 232oC as an

approximate maximum for design, we can have the opportunity to raise the temperature to

compensate for a 50% loss in activity for the LTS catalyst.

The rate data in (5-3) is based on activities of the catalyst at mid life, so it is expected that the

unit will perform better than the design at the outset. Designed outlet CO values can be kept

by adopting the strategy of increasing temperature over a long period of time. This is done in

view of the fact that activity of catalyst will decline towards the last half of its life. Increasing

temperature further however, can result in sintering of the catalyst sintering, leading to lower

activity of the catalyst. Hence, increasing temperature in such case does not remedy the

situation at all.

5.3.2 Pressure

The LTS converter will be operating at the pressure of 24.86 bar. Much care has to be taken

to prevent a high pressure drop across the LTS converter. This is to facilitate a high enough

pressure for the Pressure Swing Adsorption unit that follows the knock-out drum.

5.3.3 Steam to CO ratio

As suggested by Rase (1977), Steam to CO ratio should be in the range of 4:1 and above.

Economic analysis on design studies has to be done before the optimum value can be

obtained. The disadvantages with using a higher steam rate will be that of higher flow rates

and larger diameter equipment.



For the HTS unit, inlet’s molar flow of H2O and CO is as follows:

Species Molar flow (kmol/hr)

H2O 6326.9

CO 1265.6

The Steam to CO ratio employed for the HTS case is (6326.9/1265.6) = 5

For the LTS unit, inlet’s molar flow of H2O and CO is as follows:

Species Molar flow (kmol/hr)

H2O 5374.4

CO 313.1

The Steam to CO ratio for this case is (5374.4/313.1) = 17.2

Explanation and justification for the high ratio used for LTS:

Both CO and H2O react to form CO2 and H2. So it is understood that both the molar flow of

CO and H2O will decrease once they go through the first water-gas-shift reaction in the LTS

converter. However, CO gets converted to a much greater extent (from 1265.6kmol/h to

313.1kmol/h) as compared to H2O. Given a much greater change in the denominator of

(H2O/CO), we can hence expect an inevitable high Steam to CO ratio for the inlet feed of the

LTS converter.

Drawing steam from the feed to the LTS converter just to meet a H2O:CO ratio of around 4:1

will be an unwise move as this will give rise to greater need for extra units prior to feeding

into the LTS converter to condense the steam and then remove the liquid phase. Higher steam

does have the advantage of the driving the water-gas-shift equation to the right, producing

more of the products. However as mentioned, the drawback of higher steam content will be

bigger equipment needed to accommodate a higher volumetric flow.



5.3.4 Design Procedure for LTS outlet compositions and Mass of Catalyst used

The design procedure suggested by Rase(1977) has been adopted and is as follows:

Component Mole Balance

∆∆∆∆W(-rCO) = (-∆∆∆∆FCO) (5-47)

FCOi+1 = FCOi – (-∆∆∆∆FCO) (5-48)

FCO2i+1 = FCO2i + ∆∆∆∆FCO2 (5-49)

Where i designates increment number

Heat Balance

Heat of reaction is based on the known inlet temperature of the increment

∑∑∑∑FjCpj(Ti+1 – Ti) = (-rCO)(- ∆ ∆ ∆ ∆HCO)Ti∆∆∆∆W = (-∆∆∆∆FCO)(- ∆∆∆∆HCO)Ti (5-50)

Algorithm principle

Basis: An ∆W of 200lb can never cause a ∆T greater than 1o, as pointed out by Rase(1977).

An increment size of 1o is then selected and the change in molar quantity (∆X) of a specie i.e

CO can be calculated for each increment. Algorithm is run with using n steps(cycles), the

number of steps required to reach target amount of CO in the outlet. With every increment

step involving 200lb of catalyst, total amount of catalyst can then calculated using n x 200lbs

The following algorithm has been set up:

1. Calculate (-rCO) at inlet conditions to increment, i.

2. Calculate (-rCO)avg = (-rCO)i + [(-rCO)i – (-rCO)i-1]/2 (skip for i = 0)

3. Calculate new flow rates: Fi+1 = Fi ± ( −ro) ∆W

4. Calculate Cp and (-∆HCO) @ Ti

5. Calculate ∆T from (5-50)

6. Ti+1 = Ti + ∆T

7. yi+1 = Fi+1/(FT)i

The amount of catalyst, including outlet compositions have been worked out with this

algorithm in MATLAB. Relevant codes for this portion have been attached in Appendix A.



5.3.5 LTS outlet compositions and Mass of Catalysts used

Results of our previous algorithm are as follows:

Mole Fraction

Outlet kmol/hr

CO 0.00338 53.41

H2O 0.32347 5114.12

CO2 0.12660 2001.53

H2 0.51192 8093.44

C2H6 0.00066 10.40

CH4 0.03200 505.87

N2 0.00198 31.22

Mass flowrate in kmol/hr 15810

Mass of LTS catalyst used(kg) 34720

Outlet T (celsius) 239.3

The water shift reaction occurring in the low temperature reactor changes the composition of

the syngas species and temperature of the syngas. The CO conversion efficiency ξcan be

used to show how much CO is converted into CO2 in the LTS converter.

(4-51)

ξfor our case = (313.10 – 53.41)/313.10

= 82.9%

Mol fraction of CO in dry basis = 0.5%

Fig.5 Molar compositions of inlet and outlet feed



Comparison of HYSYS and MATLAB results for outlet feed molar compositions

Species HYSYS MATLAB

CO 0.003390 0.003378

H2O 0.323482 0.323474

CO2 0.126590 0.126599

H2 0.511910 0.511919

N2 0.000658 0.000658

CH4 0.031999 0.031997

C2H6 0.001980 0.001975

As we can see from the table above, the molar compositions for HYSYS agree well with

MATLAB‘s. Conversion value of 82.9% for CO, previously calculated in the MATLAB

environment, is inputted as part of the specifications in the Conversion Reactor module under

HYSYS.

5.3.6 Design Procedure for Aspect Ratio

As an initial approach to calculating the dimension of the catalyst bed, the aspect ratio is first

derived. The aspect ratio is defined as follows:

Aspect Ratio (AR) = (Length of Bed/Diameter of bed) (4-52)

The following equations are used:

(4-53)

(4-54)

(4-55)

(4-56)



(4-57)

The MATLAB codes for this portion are located under Appendix A2.

Algorithm principle

For every loop in the algorithm, there is a small increment of bed diameter accompanied by a

very small increment in bed height involved. Both variables are linked based on the volume

of the catalyst used for the design. Corresponding pressure drop, followed by the aspect ratio

are generated for every cycle.

In this case, the controlled term is pressure drop, which we have fixed to be at 1 bar max.

Generating a graph displaying Aspect Ratio against Pressure-Drop, we read off the aspect

ratio which corresponds to a pressure of 1 bar-drop.

Justification for controlling pressure drop at 1 bar:

Industrially the pressure drop across the LTS converter is no more than 1 bar. A large

pressure drop is uneconomical process-wise due to the usage of a pressure-swing adsorption

unit in the downstream. Furthermore, too high a pressure drop can damage the expensive

catalyst used for LTS.

5.3.7 Results for Aspect Ratio

Fig.6 Graphs of Aspect Ratio against Pressure Drop



The aspect ratio which corresponds to a pressure drop of 1 bar is 1.2287.

Using this aspect ratio, the length and height of the catalyst bed can be determined in the next

section

5.3.8 Design Procedure for the dimensions of bed and thickness of vessel wall

The MATLAB codes for this portion are located under Appendix A3.

Formulas used are as follows:

Volume of bed = (Mass of catalyst used * /90) (4-58)

By using ‘90’, the bulk density of the LTS catalyst, Equation (4-58) has already taken in

account space occupied by the voids in the catalyst packing

Diameter of bed = (4 * Volume/( ππππ * Aspect ratio))1/3 (4-59)

Length of bed = Diameter of bed * Aspect ratio (4-60)

Thickness of vessel wall

= Pressure(psi) * (Diameter/2)/(Allowable stress value*1000 – 0.6*Pressure) (4-61)

Allowable stress value is based on the material that has been used for the construction of

vessel.

For this design, low-alloy steel A387 Gr.12 has been used for the construction of the LTS

converter given its resistance to hydrogen which is existent in the process stream. The

allowable stress value for this material is given as 13.8 Kips/in2



5.3.9 Results for bed dimensions and wall thickness

Bulk Volume of Catalyst (m3) 848.71

Diameter of Catalyst Bed (m) 2.92

Length of Catalyst Bed (m) 3.59

Wall thickness (in) 0.13

For this design, a guard bed will not be considered in view of the fact there is no sulfur

content at all in our process streams. Even if there is, I am making the assumption in this case

that all sulfur poisons will be adsorbed on the catalyst in the HTS converter, hence there is no

chance for any sulfur contents to come into contact with the LTS converter.

In reality however, due to the high sensitivity of the low temperature catalyst to sulfur

poisoning, a guard bed containing zinc oxide or other guard solids is usually positioned on

top of the catalyst bed. The guideline given by Rase(1977) is that the height for a good

distribution of the catalyst will be at 100 times the diameter of the catalyst particle.



5.3.10 Allowances set for design

• Temperature

As highlighted in a previous section, I have set a design temperature of 232oC

suggested by Rase(1977), approximately 12oC higher than the intended value of

220oC, with the assumption that this will be able to compensate for 50% loss in

activity for the catalyst

• Amount of catalyst used

The amount of catalyst used will be designed with 10% extra. No references have

been found addressing this portion on the ‘extra amount of catalyst’ to use. I have

chosen a reasonable figure of 10%-extra, with the priority of bringing outlet CO

concentration down to the desired level of less than 0.5% (dry basis) in mind. Using

excessive catalyst give rise to higher catalyst cost as well as a larger reactor to

accommodate the larger volume.

New catalyst amount = 34720 * 110% = 38192kg

• Size of catalyst bed and wall thickness

Incorporating the new catalyst amount into the calculations, corresponding

dimensions for the new catalyst bed are 3.01m(diameter) * 3.70m(length)

Again, a 10% extra allowance is made for the size of bed and wall thickness.

Before After

Bed Diameter (m) 3.01 3.31

Bed Length (m) 3.70 4.07

Wall Thickness (in) 0.13 0.14

Any space that is not taken up by the catalyst, in case of an over-design, can be easily

filled up by the inert ceramic balls which are used to support the catalyst.



5.3.11 Study of controlling parameters

Pressure drop vs bed length

It is observed pressure drop increases exponentially with an increase in bed length. As

mentioned in a previous section, maximum pressure drop has been controlled at a value of 1

bar since industrially pressure drop across the LTS converters are usually not more than 1 bar.

It is evident from this graph that increasing bed length can result in an increasing increment

of pressure drop each time. This means that an infinitely long catalytic reactor can result in a

plunge of pressure drop to atmospheric pressure. Hence there is a need to control the total

amount of catalyst use as more catalyst will call for the need of larger beds.



5.4 CHOICE OF A REACTOR BED

Fig.7 Single Adiabatic bed incorporating measurements for bed size

Since the water-gas-shift is only moderately exothermic, the single adiabatic bed will be a

suitable reactor for performing the low temperature shift.

The feed will flow downwards from the top through a bed of the LTS catalyst packed in

between layers of inert ceramic support balls, which serve in creating even flow distribution

over the entire cross section and for separating solid contaminants that may have entered the

feed. The ceramic balls in the lower region act as a support while the ones at the top region

prevent movement of catalyst particles by high-velocity gas. Alternatively, a metal grid with

holes sized somewhat smaller than this catalyst can be used in place of the ceramic balls.

Rase has pointed out that in cases where flow rates are large and pressure drop must be

minimized, large diameter short beds will come into the picture. This seems to be the case for

our design.



5.4.1 Cost estimation for the LTS converter

Material Cost for Converter

Calculation of the amount of material i.e. low alloy steel A387 has been simplified to obtain

an approximate value. Assuming two half-spherical ends (at both sides) with diameter 3.31m,

total surface area of the reactor can be calculated as follows:

Surface area of reactor = π D * L + 4π(D/2)2

= π ∗ 3.31 * 4.07 + 4π(3.31/2)2

= 76.74 m2

Given a design thickness of 0.14in (0.0035m), total volume of material required

= 76.74 * 0.0035

= 0.26859 m3

Density of the steel is around 8g/cm3 = 8000kg/m3

Hence, mass of the steel required = 0.26859 * 8000

= 2148.72kg

Cost of low alloy steel (Cr-Mo) is quoted as around USD$700 – 850/T, as stated in the book

by Coulson and Richardson. For design purpose, we will take the higher value of USD$850/T.

Hence cost of the steel required = USD$850 * (2148.72/1000)

= USD$1826.41

This calculation fails to take into account material required for additional features of the

reactor, e.g. the flanges. Hence the amount of steel, and consequently the cost of steel used

could have been much more.

An important assumption made in the calculation of this material cost is an uniform wall

thickness of 0.14inch for the reactor. This may not hold in reality as thickness can be slightly

larger in areas where a greater stress is experienced i.e. bottom of reactor to handle the weight

of the catalyst and inert ceramic balls. Cost of the inert ceramic balls will not be included in

the cost calculation. Due to the inavailability of its cost info online, another assumption is

made: Cost of the ceramic balls is assumed a total cost of not more than USD$500. In my

view, this is a fair assumption as it is well known that the LTS catalyst accounts for a

significantly large portion of the LTS converter’s total cost.



Catalyst cost for converter

Copper-Zinc oxide has been quoted by Rase(1977) to be approximately USD$75/ft3. Since

total bulk volume of the catalyst in this case is (84022/90) = 936ft3,

Cost of the catalyst = 936 * 75

= USD$70,200

Factoring in the inflation rate

It ought to be noted that the cost figure quoted by Rase for Copper-Zinc Oxide was 30 years.

Over these 30 years, an increase in price has in fact taken place. The same goes for the price

of the low alloy steel A387 which was last quoted back in 2005. Hence the cost for both the

catalyst and the steel required has to be adjusted to get us a better approximation to the actual

situation. The Consumer Price Index could be a good reference to account for the amount

inflation to be factored in.

Fig.8 CPI for USA from 1913-2006

An index of 100 (base case) has been assigned to the year of 1980. From year 1977 to 2008,

it is approximately a jump from index 65 to 210 in 2008 (Value for 2008 has been

approximated from the trend in the graph). From 2005 to 2008, it is probably an increment of

10 index, from 190 to 210. The adjusted total cost of the reactor is as follows:

Cost of LTS Converter(adjusted) = USD$70,200 * (210/65) + 1826.41 * (210/190) + 500

= USD$229,318



5.5 DESIGN OF THE KNOCK-OUT DRUM

Fig.9 Knock-out drum in the HYSYS environment

5.5.1 Working principle of the knock-out drum

The knock-out drum is basically a vapor-liquid separator drum which is in the form of a

vertical vessel into which a liquid and vapor mixture (or a flashing liquid) is fed. Within the

drum, the liquid is falls to the bottom of the vessel because of gravity and hence a separation

is achieved, the liquid phase(water for this case) is then withdrawn. The vapor, on the other

hand, travels upward at a design velocity which minimizes the entrainment of any liquid

droplets in the vapor as it exits the top of the vessel.

5.5.2 Sizing of the knock-out drum

The size of the knock-out drum will be dictated by the anticipated flow rate of vapor and

liquid from the drum. Based on the assumption that those flow rates are known, the following

sizing methodology is proposed:

(1) A vertical pressure vessel with a length-to-diameter ratio of about 3 to 4 is used, vessel is

sized to provide about 5 minutes of liquid inventory between the normal liquid level and the

bottom of the vessel (with the normal liquid level being at about the vessel's half-full level)



(2) The vessel diameter is calculated using the Souders-Brown equation to determine the

maximum allowable vapor velocity:

(5-62)

Where V=maximum allowable vapor velocity, m/s

ρL= liquid density, kg/m³

ρV= vapor density, kg/m³

k= 0.107 m/s (when the drum includes a de-entraining mesh pad

Then the cross-sectional area of the drum (A) is obtained from:

A, in m² = (vapor flow rate, in m³/s) ÷ (vapor velocity V, in m/s) (5-63)

And the drum diameter (D) is:

D, in m = [(4) (A) ÷ (3.1416) ] 0. 5 (5-64)

5.5.3 Results and cost estimation

The following values have been obtained from HYSYS

ρL = 993.3 kg/m3

ρV = 9.92kg/m3

Vapor mass rate = 32.17kg/s

Vapor volumetric rate = 32.17/9.92

= 3.24m3/s

V = (0.107) ((993.3-9.92)/9.92))1/2

= 1.07m/s

Cross sectional area (A) of drum = 3.24/1.07

= 3.03



Diameter (D) of drum = [(4 * 3.03)/3.1416]0.5

= 1.96m

Height of drum = 1.96 * 4

= 7.84m

Treating shape the knock-out drum as a cylinder with a thickness similar to that of LTS

converter,

Surface area of the knock out drum = 2 * π (1.96/2)2 + π*1.96*7.84

= 54.49m2

Low-alloy steel A387 Gr.12 has been used for the construction of the knock-out drum

Assuming a thickness of 0.14in,

Volume of A387 used

= 54.49 * 0.025 * 0.14 = 0.19 m3

Mass of A387 required

= 0.19 * 8000 = 1520kg

Cost of the A387

= USD$850 * (1520/1000) = USD$1292

Cost of the A387(adjusted by CPI)

= USD$1292 * (210/190) = USD$1428

Cost of the De-entrainment mesh pad, inlet diffuser and liquid level control valve have to be

included as well. Due to the inavailability of their price, a value not more than US$2000 is

assumed in this case for the total cost of these equipment.

Hence a total of USD$(1428+2000) = USD$3428 is approximated for the knockout drum.



LITERATURE REVIEW

Developments in Water Gas Shift

Catalysis using Fe-Cr and CuO/ZnO for the HTS and LTS converters has not changed much

for the past 40 years, but continue to be improved marginally. The efforts for improving the

water gas shift reactor performance have been focused on a wide spectrum of subjects. This

involves the modification of the conventionally used catalysts improve overall activity and

stability. In particular, researchers have been making attempts at in developing catalysts

which are more tolerant to sulfur contents.

In comparison with the conventional copper-zinc water-gas-shift catalyst, researchers1 have

reported that noble metal (such as gold, silver, platinum, palladium and rhodium) catalysts

have the advantage of high activity and eliminating the self-heating issue. Langmuir-

Hinshelwood (LH) kinetics have been used to derive the rate equations involving the use of

these newly-developed metal catalysts. A good fit with the experimental data has been

obtained and it is suggested that the LH kinetic model could be a suitable one for application

in catalysis of WGS by other metals. In other papers2, theoretical studies have been

performed to formulate new kinetic expressions for the novel catalysts, with the goal of a

better control over the water-gas-shift’s reaction rate and conversions.

Still, up to now researchers have been trying to optimize the industrial water-gas-shift

reaction by tuning the various parameters involved in the process. Model-based reactor

optimization has been employed for various reactor configurations such as microreactors3,

monolith reactors4 and membrane reactors5 with the objective to achieve an overall reduced

volume for the reactors or a better energy-integration within the system. Furthermore, various

reactor configurations are analyzed in order to find out the limiting values of the main design

variables. In a work by Javier & Co-workers6, it is found that insulating material of the

reactor plays a major role in the shift converters in the sense there exists an optimal thickness

of the insulator that affects the final volume of the reactor as well as other design variables.

Such results from this study will be useful for estimating the minimum and relative sizes that

allows conventional reactor technology.

1 Jian Sun, Joel DesJardins, John Buglass, Ke Liu “Noble metalwater gas shift catalysis: Kinetics study and reactor design 2 M. Levent, Int. J. Hydrogen Energy 26 (2001) 551–558 3. G. Kim, J.R. Mayor, J. Ni, Chem. Eng. J. 110 (2005) 1–10. 4. A.S. Quiney, G. Germani, Y. Schuurman, J. Power Sources 160 (2006) 5. A. Brunetti, Barbieri, E. Drioli, K.-H. Lee, B. Sea, D.-W. Lee, Chem. Eng.Process. 46 (2007) 119–126

6 Javier A. Francesconi, Miguel C. Mussati, Pio A. Aguirre “Analysis of design variables for water-gas-shift reactors



Recently a new application for the water gas shift is in the reforming systems for fuel cells.

Fuel cell development has seen remarkable progress in the past decades because of an

increasing need for enhanced energy conversion efficiency and because of serious concerns

about the environmental consequences of using fossil fuels for electricity production.

The water gas shift reaction in fuel cell application has been studied extensively to obtain

highly accurate kinetics expressions in order to create a tool for an integrated and optimized

simulation of a whole fuel processing system. In addition, the new application of hydrogen

gas as a raw material for fuel cells for mobile power sources (PEM fuel cells) requires that

the anode inlet gas have a CO concentration lower than 10-20ppm. Otherwise, the anode is

poisoned and the cell efficiency will drop abruptly. This explains why a water gas shift has to

be employed not just to produce the hydrogen fuel but also to reduce the CO concentration.

In a work by Zalc7, simulation of a fixed bed reactor was carried out, and the water gas shift

reaction forms part of a purification train for a 10kW PEM fuel cell. In that work, a

commercial Cu/Zn/Al2O3 catalyst doped with Ba was used because it showed a higher

activity than the traditional one. A one-dimensional heterogeneous model was applied in that

simulation, and a parametric sensitivity analysis was carried out for some of the process

variables, with the purpose of finding criteria to minimize the reactor volume.

As we can see from the examples above, extensive research has been made into the water gas

shift reaction, with a significant movement towards developing the fuel cell technology. The

water gas shift reaction has played a significant role for the last 60 years at least,

conventionally employed for the production of hydrogen gas for ammonia synthesis. And it

can very well be one of the most important reactions for the 21st century, especially when

fuel cell technology has overtaken all other forms of energy production to claim leader in

fuelling our future. At present, all research made into the water gas shift is definitely a

worthwhile investment.

7 J. Zalc, V. Sokolovskii, D. Loffler, J. Catal. 206 (2002) 169–171



CONCLUSION

LTS converter specifications and cost summary

Inlet Composition Outlet Composition

CO 0.0198 0.0034

H2O 0.3399 0.3235

CO2 0.1102 0.1266

H2 0.4955 0.5119

N2 0.0007 0.0007

CH4 0.0320 0.0320

C2H6 0.0020 0.0020

Diameter Height

Dimensions 3.31 4.07

Mass of Catalyst used 34720kg

Inlet Temperature 220 [C]

Outlet Temperature 239.3 [C]

Total Cost USD$229,318

H2 yield at the exit of the LTS : 75.6mol%(Dry Basis)

Knock-out drum specifications and cost summary

Diameter Height

Dimensions 3.31 4.07

Total Cost USD$3,428

In this report, the detailed design of a fixed bed catalytic reactor for the low temperature shift

was presented. Based on the energy and mass balances, the fixed bed catalytic reactor was

modeled in MATLAB. All specifications and costing are presented in the tables above. An

study made on the pressure drop along the bed length shows the exponential relationship of

increasing pressure drop with increasing bed length.

Due to the inavailability of the price data for equipments such as the de-entrainment mesh

pad, inlet diffuser, liquid level control valve and the inert ceramic balls, conservative price

values have been assumed for these equipment and incorporated in the cost mode l of our

design. It is important to take into account the assumptions that have been established for our

modeling. Overall, design and operational optimizations have been completed for this work.



BIBLIOGRAPHY

1. Chemical reactor design for process plants / Howard F. Rase

2. Fixed-bed reactor design and diagnostics : gas-phase reactions / Howard F. Rase

3. Handbook of commercial catalysts : heterogeneous catalysts / Howard F. Rase.

4. Perry's chemical engineers' handbook

5. Modelling, simulation, and optimization of industrial fixed bed catalytic reactors

/ S.S.E. H. Elnashaie and S.S. Elshishini.

6. http://en.wikipedia.org/wiki/Souders-Brown_equation



APPENDIX A1

Matlab codes

• Calculation of Catalyst Weight and Outlet Compositions

T= input('Please enter initial temperature for reactor in K:');

F= input('Please enter flowrate for reactor in kmol per hour:');

yco= input('Please enter initial mol frac of CO:');

yh2o= input('Please enter initial mol frac of H2O:');

yco2= input('Please enter initial mol frac of CO2:');

yh2= input('Please enter initial mol frac of H2:');

yc2h6= input('Please enter initial mol frac of C2H6:');

ych4= input('Please enter initial mol frac of CH4:');

yn2= input('Please enter initial mol frac of N2:');

Fco= F*yco;

Fh2o= F*yh2o;

Fco2= F*yco2;

Fh2= F*yh2;

Fc2h6= F*yc2h6;

Fch4= F*ych4;

Fn2=F*yn2;

k=exp(12.88-1855.6/T);

K=exp(-4.72+4800/T);

rate=4*k*(yco*yh2o-yco2*yh2/K)/(379*90);

Fco1= Fco -rate*200;

Fh2o1= Fh2o-rate*200;

Fco21= Fco2 + rate*200;

Fh21= Fh2 +rate*200;

cpco= (3.376+(0.557/10^3)*T -(0.031*10^5)*T^(-2))*8.314;

cpc2h6=(1.131+(19.225/10^3)*T-(5.561/10^6)*T^2)*8.314;

cph2o=(3.47+(1.45/10^3)*T +(0.121*10^5)*T^(-2))*8.314;

cpco2=(5.457+(1.045/10^3)*T -(1.157*10^5)*T^(-2))*8.314;

cph2=(3.249+(0.422/10^3)*T +(0.083*10^5)*T^(-2))*8.314;

cpn2=(3.28+(0.593/10^3)*T +(0.04*10^5)*T^(-2))*8.314;

cpch4=(1.702+(9.081/10^3)*T -(2.164/10^6)*T^2)*8.314;

deltaa = 5.457 + 3.249 - 3.376 - 3.470;

deltab = (1.045 + 0.422 - 0.457 - 1.450) * 10^-3;

deltad = (-1.157 + 0.083 - (-0.031) - 0.121) * 10^5;

integral =(deltaa * 298.15 * ((T/298.15) - 1) + (deltab/2)*(298.15^2)*(((T/298.15)^2)-1) +

(deltad/298.15)*(((T/298.15) - 1)/(T/298.15)));

dHco = -41166 + 8.314 * integral;

dT = ((rate * 200) * (-

dHco))/(Fco*cpco+Fh2o*cph2o+Fco2*cpco2+Fh2*cph2+Fc2h6*cpc2h6+Fn2*cpn2+Fch4*cpch4);

T1 = T + dT;

yco1 = Fco1/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

yh2o1= Fh2o1/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

yco21= Fco21/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

yh21= Fh21/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

ydryco= (yco1/(1-yh2o1));

i=1;

while ydryco>0.005,

k=exp(12.88-1855.6/T);

K=exp(-4.72+4800/T);

rate=4*k*(yco*yh2o-yco2*yh2/K)/(379*90);

k=exp(12.88-1855.6/T1);



K=exp(-4.72+4800/T1);

rate1=4*k*(yco1*yh2o1-yco21*yh21/K)/(379*90);

rateavg= (rate+rate1)/2;

Fco = Fco1;

Fh2o = Fh2o1;

Fco2 = Fco21;

Fh2= Fh21;

Fco1= Fco -rateavg*200;

Fh2o1= Fh2o-rateavg*200;

Fco21= Fco2 + rateavg*200;

Fh21= Fh2 +rateavg*200;

cpco= (3.376+(0.557/10^3)*T1 -(0.031*10^5)*T1^(-2))*8.314;

cpc2h6=(1.131+(19.225/10^3)*T1-(5.561/10^6)*T1^2)*8.314;

cph2o=(3.47+(1.45/10^3)*T1 +(0.121*10^5)*T1^(-2))*8.314;

cpco2=(5.457+(1.045/10^3)*T1 -(1.157*10^5)*T1^(-2))*8.314;

cph2=(3.249+(0.422/10^3)*T1 +(0.083*10^5)*T1^(-2))*8.314;

cpn2=(3.28+(0.593/10^3)*T1 +(0.04*10^5)*T1^(-2))*8.314;

cpch4=(1.702+(9.081/10^3)*T1 -(2.164/10^6)*T1^2)*8.314;

integral = (deltaa * 298.15 * ((T1/298.15) - 1) + deltab/2*298.15^2*((T1/298.15)^2-1) +

deltad/298.15*(((T1/298.15) - 1)/(T1/298.15)));

dT = ((rateavg * 200) * (-

dHco))/(Fco*cpco+Fh2o*cph2o+Fco2*cpco2+Fh2*cph2+Fc2h6*cpc2h6+Fn2*cpn2+Fch4*cpch4);

yco = yco1;

yh2o = yh2o1;

yco2 = yco21;

yh2 = yh21;

yco1 = Fco1/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

yh2o1= Fh2o1/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

yco21= Fco21/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

yh21= Fh21/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

ydryco= yco1/(1-yh2o1);

x(i) = i*200/2.2;

y(i) = T;

a(i) = ydryco;

T = T1;

T1 = T + dT;

i= i + 1;

end

yc2h6 = Fc2h6/(Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

ych4 = Fch4/ (Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

yn2 = Fn2/ (Fco+Fh2o+Fco2+Fh2+Fn2+Fch4+Fc2h6);

Wt=i*200/2.2;

fprintf('mass of catalyst is %f kg.\n',Wt);

fprintf('outlet temperature of LTS is %f K.\n',T1);

fprintf('Outlet CO mole ratio is %f \n',yco1);

fprintf('Outlet H2O mole ratio is %f \n',yh2o1);

fprintf('Outlet CO2 mole ratio is %f \n',yco21);

fprintf('Outlet H2 mole ratio is %f \n',yh21);

fprintf('Outlet C2H6 mole ratio is %f \n',yc2h6);

fprintf('Outlet CH4 mole ratio is %f \n',ych4);

fprintf('Outlet N2 mole ratio is %f \n',yn2);

fprintf('Outlet CO amount is %f \n',Fco1);

fprintf('Outlet H2O amount is %f mol.\n',Fh2o1);

fprintf('Outlet CO2 amount is %f mol.\n',Fco21);

fprintf('Outlet H2 amount is %f mol.\n',Fh21);

fprintf('Outlet C2H6 amount is %f mol.\n',Fc2h6);

fprintf('Outlet CH4 amount is %f mol.\n',Fch4);

fprintf('Outlet N2 amount is %f mol.\n',Fn2);

plot (x,y);

plot (x,a);



APPENDIX A2

• Determination of Aspect Ratio

mass = input('Please enter mass of catalyst in lbs:');

flowrate = input('Please enter flowrate in lbs per hour:');

u1 = input('Please enter input viscosity in cp:');

u2 = input('Please enter output viscosity in cp:');

pf = input('Please enter density of feed in lbs per feet3:');

u = (u1+u2)/2;

V = mass/90;

for i=1:600;

hc = i*0.1;

dc = 2*((V/(pi*hc))^0.5);

AR = hc/dc;

area = pi*(dc/2)^2;

G = flowrate/area;

Dp = 0.15/12;

Nre = Dp*G/(2.42*u);

fk = 1.75+150*(0.555)/Nre;

dPft = hc*((fk*G^2/(Dp*pf*32.17*(3600^2)))*0.555/(0.445^3));

dPsi = dPft/(12^2);

dPbar = dPsi/14.7;

x(i,1) = AR;

x(i,2) = dPft;

x(i,3) = dPsi;

x(i,4) = dPbar;

x(i,5) = G;

% if dPsi == 14.7,

% fprintf ('AR = %f.\n',x(i,1))

% end

end

plot(x(:,4),x(:,1));

APPENDIX A3

• Determination of dimensions of catalyst bed and wall thickness

mass = input('Please enter mass of catalyst in lbs:');

AR= input('Please enter AR:');

P = input('Please enter pressure in psi:');

S = input('Please enter maximum allowable stress value in kips per inch square:');

V = mass*144*12/90;

E = 1;

D = (4*V/(pi*AR))^(1/3);

L = D*AR;

t=P*(D/2)/(S*1000*E-0.6*P);

fprintf ('t= %f inches \n',t);

fprintf ('Bulk Volume of catalyst = %f cubic feet \n',V);

fprintf ('Diameter of Catalyst Bed = % f ft \n',D);

fprintf ('Length of Catalyst Bed = %f ft \n',L);

CN 4120: Design II Team 32: Zhang Zihong (U046816H) PSA Unit Design Report


Chapter 6 : PRESSURE SWING ABSORPTION

6.1 INTRODUCTION

Conventionally, the industrial purification of hydrogen from steam-methane reforming will

involve the use of a CO2 scrubber with aqueous monoethanolamine (MEA) and a methanator

to convert the unreacted CO to CH4 before recycling it back to the Steam-methane reformer

as feed. However, this purification method gives rise to the problems of personnel safety and

solvent disposal due to the toxicity of MEA [1].

Pressure Swing Adsorption (PSA) has been increasingly adopted as the preferred mode of

purification in the production of oxygen and hydrogen due to the advantages it possessed

over its rivals in terms of selectivity, throughput and efficiency. Compared to the CO2

scrubber (95 - 97% in product purity), the attainable product purity for a typical PSA system

is 99.99% [2]. This can result in a possible improvement in the refinery downstream

operating margin, since a purer treat gas to the hydrotreaters can minimize the effect of

catalyst coking and deactivation, which translates to a lower operating cost. Furthermore,

PSA would have eradicated the problem of solvent disposal since it is a solvent-free process.

However, due to its high operating pressures of up to 20 bar, additional reinforcement of the

process equipments is required which might lead to a higher initial capital cost.

In this report, the PSA configuration is incorporated in the design of a hydrogen plant that is

situated in Singapore. The PSA’s feed is directed from the knockout drum that is downstream

of the low temperature shift reactor (LTS) and the stream composition is approximately in the

range of 75% H2, with the remainder comprising of CO2, CO, CH4 trace amounts of N2 and

H2O.




In the interim report that was submitted earlier, a polybed PSA system of 7 – 10 beds was

proposed to recover up to 85% of the hydrogen product at a purity of 99.9%. According to

the Hysys simulation generated for the interim report, table 1 is a summary of the

composition of the incoming stream from the knockout drum to the PSA:

Table 1: Hysys screenshot of the composition of feed entering PSA

Thus the objective of this report is to design a PSA system which is able to achieve a

production of 1.25 x 109 m3 (STP)/yr at a product purity specification of 99.9%.

6.3 THEORETICAL BACKGROUND

6.3.1 Separation via adsorption

Adsorption is a process in which there is a selective transfer of solutes (adsorbates) in the

fluid phase to the surface of solid particles (adsorbents) through the formation chemical

bonds or electrostatic attractions [3]. Figure 1 illustrates the accumulation of adsorbates such

as CO2 on the internal pores surfaces of adsorbents, such as activated carbon for which a

highly porous structure is required to achieve a large surface area for adsorption per unit

volume of adsorbents used.



Figure 1: Diagram on adsorption [4]

6.3.2 Pressure-Swing Adsorption (PSA)

A pressure-swing adsorption is a process that selectively separates certain gaseous

components from a gas mixture by effecting a change in the system pressure. The gas

components can be separated either via their molecular characteristics or affinity for an

adsorbent material. Under high pressure, the adsorbates are selectively adsorbed onto the

adsorbents, which is the crux of the whole separation process. The system pressure is

subsequently lowered so as to effect the desorption of the adsorbates, thereby allowing the

regeneration of the bed.

6.3.3 Skarstrom Cycle

The design of this report’s PSA system is based on the basic form of Skarstrom Cycle [5]

which comprises of 2 beds and 4 basic steps:

• Pressurization

o Pressurization of the PSA unit increases the affinity of the adsorbate with the

absorbent bed which causes enrichment of the less selectively adsorbed species.

This is done using the feed.



• High Pressure Adsorption

o The raffinate is recovered during this process until the absorbent bed is relatively

saturated.

• Blow-down

o Depressurization of the PSA unit occurs at this step, which allows the adsorbed

species to desorb from the bed and be removed.

• Purge

o The bed is purged with part of the raffinate from the HPA step of another bed.

This is to regenerate the bed and allow it to be used for another cycle.

6.3.4 Adsorbents

Figure 2: Picture of activated carbon

Activated carbon, is a general term that includes carbon material derived mostly from

charcoal. Possessing an exceptionally high surface area due to a high degree of microporosity,

it is used in the PSA process for the adsorption of CO2 and CH4.



Zeolites are hydrated aluminosilicate minerals commonly referred to as "molecular sieves".

The term molecular sieve refers to a particular property of these materials with the ability to

selectively sort molecules based primarily on a size exclusion process. This is due to a very

regular pore structure of molecular dimensions. Zeolites are considered crystalline while

activated carbon is amorphous. Zelites are used in PSA design to remove N2 and CO from the

effluent gas.

6.4 DESIGN CONSIDERATIONS

Being a dynamic process, the model for PSA involves the evaluation of nonlinear partial

differential equations. The high level coupling of the model parameters further adds to the

complexity of the whole simulation process. Thus before the commencement of the

simulation process, an appropriate model should be selected and certain assumptions would

have to be adopted so as to reduce the modeling complexity.

The assumptions made were as follows [5]:

1. An isothermal system is assumed.

2. There is negligible frictional pressure drop along the bed length.

3. Mass transfer between the gas and the adsorbed phases is accounted for all the steps

and is sufficiently represented by the Linear Driving Force (LDF) model. During the

adsorption and the purge steps, the total pressure in the bed remains constant while it

varies exponentially with time during the blow-down and pressurization.

4. Fluid velocity in the bed varies along the length of the column, as determined by the

overall mass balance.

5. The flow pattern is described by the axial dispersed plug flow model.

6. Equilibrium relationships for the components are represented by extended Langmuir

isotherms.

7. The ideal gas law applies.

8. The PSA process is assumed to be equilibrium controlled.



In the LDF model, the mass transfer rate equation was represented as:

( )iiii qqk

t

q−=

∂

∂ *

where

o

M

Mp

ip

pp

i

cq

c

D

DD

BorAiq

c

r

Dk

withmequilibriuinionconcentratphasesolid

gasfeedinionconcentratgas

3)(tortuosity

Equation)Enskogy(chapmandiffusivitmolecular

)15(parameterLDF

, where

)(

0

0

0

0

2

=

−

=Ω

=

=

Ω=

τ

τ

ε

In order to simplify the model, the mass transfer coefficient, ik for macropore control was

adopted by assuming that the adsorbents were spherical in shape with no micropore diffusion

of the adsorbates.

By neglecting the interactions between the adsorbed components and assuming that the

reduction of the vacant surface area for the adsorption of A is solely due to the adsorption of

the other components, the extended Langmuir model was applied as follows:

( )∑+

=

j

jj

imii

ipK

pqKq

1

where ( )miq is the maximum amount of adsorption of species i for coverage of the entire

surface.



6.5 ACTUAL MODELING OF PSA

Even after simplifying the LDF model, a moderate degree of difficulty still exist in the

simulation of PSA via the use of COMSOL. Coupled with the issue of time constraint, the

complexity of the model was further reduced to as follows:

1. All the water was assumed to be removed via the use of a layer of silica gel before the

gas stream entered the PSA.

2. The subsequent six component system was reduced to a binary system of hydrogen

(carrier gas) and carbon dioxide (adsorbate).

3. The molecular diffusivity of carbon dioxide in hydrogen, DM was derived using

Hirschfelder Equation for which the molecules were assumed to be non-polar and

non-reacting.

4. The pure component isotherms and the respective constants on a common adsorbent

were applied.

5. The Peclet number was assumed to be a constant due to its small numerical value.

6. The product of time constants, a1 and a2 and time, t for both pressurization and blow-

down step were assumed to be 6.

The selection of hydrogen and carbon dioxide as the two components was in accordance with

the Hysys simulation whereby they accounted for more than 93% of the total number of

moles entering PSA (CO2 ≈ 19%, H2 ≈ 75%). Thus for conservative design, the mole ratio of

H2 to CO2 was assumed to be 1:3. Activated carbon [7] was employed as the adsorbent in this

simulation due to its higher affinity for CO2 as compared to H2.



The following partial differential equations were employed:

6.5.1 Component Mass Balance

For which A = CO2 and B = H2,

0t

q1

t

c

z

vc

z

cv

z

cD AA

AA

2

A

2

L =∂

∂

εε−

+∂

∂+

∂∂

+∂

∂+

∂

∂−

6.5.2 Overall Mass Balance

0)t

q

t

q(

1

t

C

z

vC

z

Cv

z

CD BA

2

2

L =∂

∂+

∂

∂

ε

ε−+

∂

∂+

∂

∂+

∂

∂+

∂

∂−

6.5.3 Pressure terms

)steps Purging and AdsorptionFor (0

)stepdown -BlowFor ())((

)steption PressurizaFor ())((

2

1

2

1

=∂∂

−−=∂∂

−=∂∂

−

−

t

P

eaPPt

P

eaPPt

P

ta

LH

ta

LH

6.5.4 Adsorption rates

)( B

e

BB

B qqkt

q−=

∂

∂ )( A

e

AAA qqkt

q−=

∂∂



The above equations were subsequently converted to dimensionless form so as to firstly,

reduce the number of variables and complexity in the modeling and secondly, to facilitate in

the subsequent scale-up to the actual plant capacity.

6.5.5 Overall Mass Balance in Dimensionless Form

P

1P)

xx(

P

TR

1q

v B

s

A0g

SA τ∂∂

−τ∂

∂γ+

τ∂∂

ε−ε

−=χ∂

∂

6.5.6 Component Mass Balance in Dimensionless Form

0x

yx

)1y(P

TR1q

yv

y

Pe

1y BAs

AA

0g

SAA

2

A

2

A =

τ∂

∂γ+

τ∂

∂−

εε−

+χ∂

∂−

χ∂

∂=

τ∂

∂

6.5.7 Dimensionless Pressure terms

Vo

La

oH

L ev

La

PPP

τ

τ111

−

−=

∂∂

(For Pressurization step)

Vo

La

oH

L ev

La

PPP

τ

τ221

−

−−=

∂∂

(For Blow-down step)

0=∂∂τP

(For Adsorption and Purging steps)

6.5.8 Dimensionless Langmuir Adsorption Isotherms

−−

++

α=τ∂

∂A

og

HAB

og

HAA

og

HAA

A

Ax

TR

PP)y1(b

TR

PPyb1

TR

PPyb

x

−−

++

α=τ∂

∂B

og

HAB

og

HAA

og

HBB

A

Bx

TR

PP)y1(b

TR

PPyb1

TR

PPyb

x



The 5 dimensionless equations were inputted into COMSOL for simulation with the set of

boundary conditions.

6.5.9 Boundary Conditions

Pressurization Adsorption

( )

0

0

1

1

1

0000

=

=

−=−

=

=

+=−===

χ

χ

χχχχ

χ

χ

v

d

dy

yyvd

dy

Pe

A

AA

A

( )

1v

0d

dy

yyvd

dy

Pe

1

0

1A

0A0A00A

=

=χ

−=χ

−

=χ

=χ

+=χ−=χ=χ=χ

Blowdown Purge

0v

0d

dy

0d

dy

1

1A

0A

=

=χ

=χ

=χ

=χ

=χ

( ))(

1

0

1

1111

0

Gratiovelocityfeedtopurgev

yyvd

dy

Pe

d

dy

AA

A

A

=

−=−

=

=

−=+===

=

χ

χχχχ

χ

χ

χ



6.6 MODEL OPTIMIZATION

The aim of the optimization process is to maximize the recovery of the H2 product whilst

attaining the specified purity of 99.9%. For a reasonable and simplified design of the PSA

system, the following heuristics would be adhered to in this report [8]:

• A two bed system was assumed for the COMSOL-Matlab simulation.

• Bed Porosity should be maintained in the range of 0.3 – 0.5.

• Adsorption time should be close to the breakthrough time.

• Adsorption and Desorption time should be kept constant for 2 bed processes.

• Purge to feed volume ratio, G should in the range of 1.0-2.0.

• The ratio of pressurization time to adsorption time should be capped at 0.2.

• The superficial velocity entering the bed should not exceed 75% of the minimum

fluidizing velocity [5].

There are multiple decision parameters that govern the operations of PSA and this includes

the bed length, bed diameter, superficial velocity of the feed and purge ratio, G. A systematic

approach was thus adopted wherein a variation of one of the parameters would be performed

with the others held constant. For this varied variable, the optimum point would be that at

which the required purity of 99.9% was achieved with the maximum allowable recovery. By

maintaining the constant value for this variable, this optimization process was subsequently

repeated for all the other parameters.



6.6.1. Process Methodology

In this report, the following optimization sequence was adopted:

1. Initial approximation of the adsorption time from the breakthrough curve.

2. Determination of the cyclic steady state.

3. Refinement of the pressurization time.

4. Optimization of feed superficial velocity and diameter of the bed.



6.6.2 Initial approximation of the adsorption time from the breakthrough curve

According to A. Varnia [9], an industry expert, a 5-bed PSA system with the height and

diameter of approximately 8.5 m and 3.2 m respectively, can achieve a hydrogen production

of (STP) 1.25 x109 m3/yr at 99.9% product purity. From this capacity and diameter, an inlet

flow-rate of 3.33 m3/s and subsequently, a velocity of 0.41m/s, were approximated with the

Hysys simulation. Thus this initial bed height, diameter and velocity were adopted in

COMSOL to estimate the time period for breakthrough to occur (the time duration before

which the maximum allowable adsorbate concentration (CO2) in the effluent gas (H2) is

exceeded). The purge to feed volume ratio, G was also assumed to be at 2 [8].

Figure 3: Breakthrough curve of yco2 with time

Based on this initial breakthrough time, an adsorption time could be subsequently

approximated to initiate the optimization process, since the duration of the high pressure

adsorption should be theoretically equal to that in order for breakthrough to occur [8].

Breakthrough occur at T = 17 τ



From the COMSOL simulation for the adsorption step, it could be observed that the time

duration for breakthrough was about 17 τ in dimensionless time and this translates to:

s

V

BTt

o

L

130

36.0/4.0

5.8*17

*

=

=

=

Thus according to figure 7.1, 130 s was the upper limit for the varying of the subsequent

adsorption times. The purity of the product started to decline once the breakthrough time was

exceeded and in order to attain a high level of product purity, it is necessary to set the

adsorption time to as near as possible to the breakthrough time.

6.6.3 Determination of Cyclic steady state

The dynamic steps in PSA meant that no steady state would be attained. However, after a

sufficiently large number of cycles, there will be a point in time whereby the profiles

achieved by the bed at the end of a cycle will be the same at the start of the next and that is

when cyclic steady state is achieved.

In order to determine the cyclic steady state, the number of cycles, h required for each run of

simulation was varied as shown below in figure 3. Four values of h (3, 4, 10 and 14) were

used. For each run of simulation, a plot of the yco2 wave front versus bed length was

generated and subsequently, they were compared on the same graph. There was a shift in the

wave front to the right with an increase in the number of cycles. However, beyond the run for

10 cycles, a constant wave front was observed, which led to the conclusion that 10 cycles was

required for cyclic steady state to be attained.



Figure 4: Plot of varying yco2 wave front versus bed length

6.6.4 Refinement of the pressurization time

The durations of the adsorption and pressurization time are two important parameters that

affect the performances of a PSA system. According to the simulations performed by S. Jain

[8], the ratio of the pressurization time and adsorption time should be capped at a limit of 0.2

so as to achieve the best possible result for both purity and recovery. According to the earlier

discussion, the optimum adsorption time should be set near to the breakthrough time and a

value of 120 s was assumed for conservative reasons in order not to exceed the breakthrough

time. The pressuration time was subsequently varied to obtain the best result. As seen from

figure 7.3, the change in pressurization time was inversely proportional to recovery.

)

(

)

(

covRe

2

2

2

2

steptionpressurizaduringfedHofamount

feedinstepadsorptionduringusedHofamount

steppurgeinusedHofamount

adsorptionduringobtainedHofamount

ery+

−

=

Run for 3 cycles

Run for 10 and 14 cycles

Run for 5 cycles



According to the above equation used in the calculation of recovery, with an increase in the

pressurization time, more of the H2 will be fed during the pressurization step and this will

result in a decrease in recovery, which was consistent with the trend displayed in figure 7.3.

The values of Tads = 120 s and Tpr = 24 s were selected for further optimization.

Purity vs Re cove ry for varying Tpr at constant Tads =

120s

0.9976

0.9978

0.998

0.9982

0.9984

0.9986

0.9988

0.999

0.9992

0.817 0.819 0.821 0.823 0.825 0.827 0.829 0.831

Figure 5: Plot of purity versus recovery at varying Tpr (Tads = 120 s)

6.6.5 Possible optimization of feed superficial velocity and diameter of the bed

If the purity specification of 99.9% was not attained or if there was potential for further

improvement in product recovery, the last parameter for optimization would be the feed

superficial velocity. In this section, the feed velocity was adjusted while maintaining a

constant Tads and Tpr. An optimum value of 0.4 m/s for feed superficial velocity was

subsequently attained from figure 7.4. It should be noted that the calculated feed superficial

velocity should not exceed 75% of the minimum fluidizing velocity. This was because

fluidization of the fixed bed could result in the possible loss of adsorbents with the exiting

product stream. The minimum fluidizing velocity was calculated to check for the validity of

the superficial feed velocity: 1619)10116.1)(36.01(

)634.9)(4.0)(003.0(

)1(Re

5=

−=

−=

−x

vD fsp

µε

ρ

Tads = 120 s Tpr = 24 s Tads = 120 s

Tpr = 22 s



Turbulent flow occurs where 500<Re<200000

To calculate Vf, the modified Ergun equation for turbulent flow [10] was used:

78.2))634.9(

003.0)81.9)(634.9850)(3((

)(32

12

1

=−

=

−=

ρ

ρρ ps

t

gdu

Particle Density, ρp = 850 kg/m3 Fluid Density, ρf = 9.634 kg/m3

g = 9.81 m/s2 Particle Diameter, dp = 0.003 m

ε = 0.36 Kinematic Viscosity, µ = 1.180 x 10-5 Pa s

Based on the above mentioned heuristics, the calculated feed superficial velocity of 0.4 m/s

was safely within the range of less than 75% of the minimum fluidizing velocity, which was

calculated to be 2.8 m/s.

Constant Tads, Tpr (120 s, 22 s), varying V

0.99700

0.99750

0.99800

0.99850

0.99900

0.99950

1.00000

0.80000 0.81000 0.82000 0.83000 0.84000

Recovery

Pu

rity

Figure 6: Plot of purity versus recovery for different feed superficial velocity

V = 0.41 m/s V = 0.40 m/s



6.7 FINAL RESULTS AND DISCUSSIONS

Adsorption Time, Tads = 120 s

Pressurization Time, Tpr = 24 s

Superficial feed velocity, V = 0.4 m/s

Recovery = 82.0%

Purity = 99.9%

timeTotal

purgeTotaloutputTotaloductivity

−=Pr = 1.15 x 109 m3 (STP)/yr

where total time is time taken for both pressurization and high pressure adsorption to occur

Required capacity = 1.25 x 109 m3 (STP)/yr

It can be observed that the productivity achieved in this simulation was lower than the

required plant capacity. One probable reason could have been due to the higher recovery

assumed in the Hysys interim report (85%), which had subsequently resulted in a lower

quantity of feed required.

It should also noted that the separation performed in this report is on a binary system, which

fails to account for the presence of the other components, such as CO, CH4 and N2 that are

present in the actual feed stream to PSA. Thus the actual purity might have in fact been lower.



6.8 COST ESTIMATIONS

The cost estimation is done by assuming a diameter of 3.0 m and bed length of 8.5 m.

Volume of vessel, LD

V4

2

π=

3

2

60

)5.8(4

3

m=

= π

From Figure A.7, purchased cost of vessel using carbon steel,

60850 ×=o

pC = $ 51,000



Pressure factor, mtforP

DP

F vesselvesselP 0063.00063.0

00315.0)]1(6.0850[2

)1(

, >+

+−+

= pp925 [1]

Where P = pressure in barg and D = diameter of vessel

P = 25 bar = 24 barg

63.70063.0

00315.0)]124(6.0850[2

3)124(

, =+

+−+

=vesselPF

Purchased cost of vessel after pressure correction and adjusted for inflation

2001SeptinCEPCI

2007NovinCEPCIFCC vessel,p

o

pp ××=

2007NovinCEPCI = 593.6 [Chemical Engineering February 2008] (latest data available)

833,581$397

6.59363.7000,51$ =××=pC

Assuming auxiliary equipment, which include expander, control valves and safety valves

piping,

Total Cp = $ 581,833 × 1.3 = $ 756,383



Calculation of cost of activated carbon

Mass of activated carbon used = V(1- ε b)ε p

= 60 × (1-0.36) × 850

= 32,640 kg

Cost of Calgon activated carbon ≈ $5/kg [2]

Cost of activated carbon = 32,640 × 5 = $ 163,200

The real PSA bed for H2 purification is packed by not only activated carbon, but also Zeolite

which is of much higher price than activated carbon.

Total cost for a PSA adsorber = $ 581,833 + $ 163,200

= $ 745,033

This estimated cost is expected to be less than the actual price because the actual PSA bed

will use not only activated carbon but also Zeolite

For 2-PSA adsorber = $ 745,033 × 2

= $ 1,490,066

Assuming installation of PSA is 10% of total cost

Therefore, total estimated cost of installed 2 PSA bed system

= $ 1,490,066 × 1.1

= $ 1,639,073



6.9 CONCLUSION

For this report, varying degrees of success has been achieved in the simulation of the PSA

system. A productivity of 1.15 x 109 (STP) m3/yr at 99.9% product purity and 82% recovery

has been obtained from the COMSOL Matlab simulation of a two bed, binary PSA system

through the variation of pressurization time, Tpr and feed superficial velocity, V. However,

this productivity still falls short of the required capacity of the H2 plant which was fixed at

1.25 x 109 (STP) m3/yr. A reasonable explanation for this shortfall could have been attributed

to the lower feed required arising from a higher product recovery assumed initially (85%).

Industrially, the attainable recovery is within 80 – 90% at a purity of up to 99.999% with a 7 -

10 bed configuration and the use of more steps such as pressure equalization, co-current

pressurization to increase the recovery of the system. Thus it might not have been feasible to

achieve a recovery of 85% at a purity of 99.9% with a complete multi-component simulation

of a two bed system.

Due to severe time constraints, only the binary system was considered in this simulation and

this could have been expanded further into a multi-component system if more time was

allowed. Furthermore, the simulation was performed for one adsorbent (activated carbon),

which is not applicable for a multi-component system where at least two adsorbents are

required for the near complete sorption of the various components so as to achieve a product

with a very high level of purity.

Finally, a buffer tank could have been considered in the overall design of the system, so as to

deal with the possible occurences of irregular flow from the upstream units. Through the

installation of a buffer tank, irregular flow rates can be eliminated and this will ensure the

smooth operation of the PSA system. Furthermore, in times of an upstream unit upset, the

avaliability of the buffer tank can continue to sustain the PSA operations until the upstream

unit is up and running again.



6.10 NOTATIONS

Symbol Meaning Symbol Meaning

c Concentration v Superficial Velocity

q Adsorbed Phase

Concentration

(mol/kg)

Vo Interstitial Velocity

K Mass transfer

coefficient

L Bed Length (m)

DL Molecular

Diffusivity

D Bed Diameter (m)

Pe Peclet Number qas Saturated Adsorbed

phase concentration

(CO2)

T Time (s) qbs Saturated Adsorbed

phase concentration

(H2)

T Temperature (K) b Langmuir constant

Ε Void Fraction Y Gas phase

concentration

z Axial position in

adsorption bed (m)

G Purge to feed ratio

Vf Minimum

fluidization velocity

g Gravitational

acceleration (m/s2)

Dp Particle Diameter µ Dynamic Viscosity

(Ns/m)



Subscripts

Symbol Meaning Symbol Meaning

A CO2 s Saturated

B H2 H, L High, low



6.11 APPENDIX

The molecular diffisivity:

Assuming 1) negligible viscosity change over temperature 2) Feed at 50ºC

A: Carbon Dioxide B: Hydrogen

Using Equation 24 -33 [6],

DAB

BA

ABP

MMT

DΩ

+

=2

2/1

2/3 11001858.0

σ

Where 2

BA

AB

σσσ

+=

BAAB εεε =

Kergsx /1038.1 16−=κ

From Appendix K, Table K.2

3.33

190

=

=

κεκε

B

A

968.2

996.3

=

=

B

A

σ

σ

)(int88216.0 erpolateD =Ω

=ABD 3.645E-2 cm2/s = 3.645E-6 m2/s



The original axial dispersed plug flow model [] is given by

0t

q1

t

c

z

vc

z

cv

z

cD ii

ii

2

i

2

L =∂

∂

εε−

+∂

∂+

∂∂

+∂

∂+

∂

∂− where

termdispersionaxialz

cD i

L =∂

∂−

2

2

vrD7.0vd5.0D7.0D pMpML +=+=

termtransfermassconvectivez

vc

z

cv i

i =∂

∂+

∂

∂

termonaccumulatit

ci =∂

∂

termphasesolidtoadsorptiont

q1 i =∂

∂

εε−

for i = A and B, A = carbon dioxide and B = hydrogen

0t

q1

t

c

z

vc

z

cv

z

cD AA

AA

2

A

2

L =∂

∂

εε−

+∂

∂+

∂∂

+∂

∂+

∂

∂−

We will first derive the overall mass balance equation which will be used to find the velocity

profile along the bed.

The OVERALL MASS BALANCE, which is the sum of component mass balances

0)t

q

t

q(

1

t

)cc(

z

v)cc()cc(

zv)cc(

zD BABA

BABABA2

2

L =∂

∂+

∂

∂

εε−

+∂

+∂+

∂∂

+++∂∂

++∂

∂−

Ccc BA =+

0)t

q

t

q(

1

t

C

z

vC

z

Cv

z

CD BA

2

2

L =∂

∂+

∂

∂

εε−

+∂∂

+∂∂

+∂∂

+∂

∂−



The total concentration is not a function of the distance of bed

)z(fCcc BA ≠=+ and

0z

C

z

C2

2

=∂∂

=∂

∂

The overall material balance becomes

0)t

q

t

q(

1

t

C

z

vC BA =

∂

∂+

∂

∂

εε−

+∂∂

+∂∂

∂

∂+

∂

∂

ε

ε−+

∂

∂−=

∂

∂

t

q

t

q1

t

C

C

1

z

v BA

0gTR

PC =

t

P

TR

1

t

C

0g ∂∂

=∂∂

P

1

t

P)

t

q

t

q(

1

P

TR

z

v BA0g

∂∂

−∂

∂+

∂

∂

εε−

−=∂∂

OVERALL MASS BALANCE

This overall mass balance equation is used to find the velocity profile along the bed

Dimensionless form

P

1P)

xx(

P

TR

1q

v B

s

A0g

SA τ∂∂

−τ∂

∂γ+

τ∂∂

ε−ε

−=χ∂

∂D’LESS OVERALL MASS BALANCE



We will then proceed to find component mass balance

0t

q1

t

c

z

vc

z

cv

z

cD AA

AA

2

A

2

L =∂

∂

εε−

+∂

∂+

∂∂

+∂

∂+

∂

∂−

RT

Pyc A

A =

0t

q1

RT

P

t

y

z

v

RT

Py

RT

P

z

yv

RT

P

z

yD AAAA

2

A

2

L =∂

∂

εε−

+∂

∂+

∂∂

+∂

∂+

∂

∂−

0t

q

P

RT1

t

y

z

vy

z

yv

z

yD AA

AA

2

A

2

L =∂

∂

εε−

+∂

∂+

∂∂

+∂

∂+

∂

∂−

Substitute OVERALL MASS BALANCE INTO COMPONENT MASS BALANCE

P

1

t

P)

t

q

t

q(

1

P

TR

z

v BA0g

∂∂

−∂

∂+

∂

∂

εε−

−=∂∂

0t

q

P

RT1

t

y

P

1

t

P)

t

q

t

q(

1

P

TRy

z

yv

z

yD AABA0g

AA

2

A

2

L =∂

∂

εε−

+∂

∂+

∂∂

−∂

∂+

∂

∂

εε−

−+∂

∂+

∂

∂−

Simplify the above equation

0t

q

P

RT1)

t

q

t

q(

1

P

TRy

z

yv

z

yD ABA0g

AA

2

A

2

L =∂

∂

εε−

+

∂

∂+

∂

∂

εε−

−+∂

∂+

∂

∂−

Rearrange it

0t

qy

t

q)1y(

P

TR1

z

yv

z

yD

t

y BA

AA

0gA

2

A

2

LA =

∂

∂+

∂

∂−

εε−

+∂

∂−

∂

∂=

∂

∂ COMP. MASS BAL

0x

yx

)1y(P

TR1q

yv

y

Pe

1y BAs

AA

0g

SAA

2

A

2

A =

τ∂

∂γ+

τ∂

∂−

εε−

+χ∂

∂−

χ∂

∂=

τ∂

∂

D’LESS COMPONENT MASS BALANCE



6.12 CONSTANTS APPLIED IN COMSOL SIMULATION

Temperature 323K

Bed voidage 0.36

Particle density 0.85 g/m3

qco2 7.12 mol/kg Saturated adsorbed CO2

qH2 4.32 mol/kg Saturated adsorbed H2

bco2 2.54e-6 Pa-1 CO2 Langmuir constant

bh2 7.02e-8 Pa-1 H2 Langmuir constant

kco2 0.1 s-1 Mass transfer coeff CO2

kh2 1.0 s-1 Mass transfer coeff H2

High P 2.5e6 Pa

Low P 1e5 Pa

[7]



6.13 REFERENCES

[1] Retrieved from ‘www.astrochemicals.com/10136.htm’, 24/03/2008

[2] Tinall, B.M and Crews, M.A, ‘Economics of export steam for hydrogen plants,’

Hydrogen Engineering, (2003) 39

[3] J D Seader and E J Henley, “Separation Process Prinicples”, John Wiley & Sons, pp 798

(1998)

[4] Retrieved from ‘http://www.norit-americas.com/images/adsorption-image.gif’

[5] ‘Pressure Swing Adsorption’, D.M. Ruthven et al, VCH Publishers, 1994, Pg 225

[6] J.R Welty, C.E.Wicks, R.E. Wilson, G.Rorrer, Fundamentals of Momentum, Heat, and

Mass Transfer, John Wiley & Sons, Inc. (2001) 4th ed

[7] J.H Park, J.N Kim, S.H Cho, Performance Analysis of Four-Bed H2 PSA Process Using

Layered Beds, AIChE Journal (2000) Vol 46

[8] S.Jain, A.S.Moharir, P.Li, G.Wozny, Heuristic Design of pressure swing adsorption: A

preliminary study, Separation and Purification Technology 22 (2003) 25-43.

[9] V.Aspi (2008) Singapore Refinery Company, Jurong Island, Singapore, Personal

communication

[10] Retrieved from

‘http://faculty.washington.edu/finlayso/Fluidized_Bed/FBR_Fluid_Mech/fluid_bed_scroll.ht

mcosting’

[11] Turton et al, Analysis, Synthesis, and Design of Chemical Processes, 2nd edition,

Prentice Hall, 2007

[12] Retrieved from

‘http://www.apswater.com/shopdisplayproducts.asp?id=157&cat=Activated+Carbon’

CN 4120: Design II Team 32: Sin Yew Leong (U046835M) HEN Unit Design Report


Chapter 7 : HEAT EXCHANGER NETWORK

EXECUTIVE SUMMARY

Energy integration is of utmost importance in a chemical plant, where huge demands are being

placed on plant personnel to fully utilize available utilities, in order to obtain the most

economical and practical solution. To achieve optimal heat integration, it would involve the

systematic development of a heat exchange network (HEN), together with detailed designs of

each and every heat exchanger.

The primary targets set out in this problem statement would be:

To design a heat exchanger network for a hydrogen plant to be sited in Singapore

To look into possible optimization of the heat exchanger network, considering a

multitude of network variations and possible network evolution

To design one of the heat exchangers in detail after the heat exchanger network has been

finalized

Pinch analysis was the method of choice in deriving a minimum energy recovery (MER) network.

Though 2 loops were identified, various infeasibilities limited the evolution of the MER network.

The final network was eventually chosen based on the minimum total annual cost. The total

annual cost of the selected network was within 1% of the target set.

Detailed thermal sizing of the SMR feed preheater was then performed. A TEMA style AES

shell and tube exchanger was utilized for this service. Computation of the estimated area, heat

transfer coefficient, pressure drops and cost were calculated. The final design for the heat

exchanger is projected to be able to handle about 99% of its intended heat load.

ACKNOWLEDGEMENTS

This section dedicates acknowledgements to all who have helped the author by offering their

valuable insights and advices. In particular, the author would like to express gratitude to Prof.

Rangaiah for his advice. Last but not least, this work would not have been possibly done without

the coordination and teamwork in Team 32 (CN4120 – Sem 2 of AY2007/08), hence the author

would like to thank all of them for their assistance and understanding.



7.1 DESIGN METHODOLOGY OF A HEAT EXCHANGER NETWORK

Pinch technology was systematically adopted for the design and evolution of this design problem.

Substantial effort was first channeled into the collection of relevant stream data. Manual

determination and verification of these stream data properties, coupled with cost optimization

were then carried out before the actual synthesis of the heat exchanger network. Two software

programs, Hysys and HX-Net, were extensively used in this design process to facilitate data

extraction, data processing, pinch analysis and super-targeting.

7.1.1 Determination & Verification of Stream Data Properties Extracted from Hysys

Stream data properties (mass flow rates, M, supply temperature, Ts, target temperature, Tt, heat

capacity, Cp and convective heat transfer coefficient, h) extracted into HX-Net from Hysys have

to be validated first before the actual design of the heat exchanger network can proceed. Data

relevant to the design are validated using suitable equations and correlations.

7.1.1.1 Calculations of Maximum Design Velocities

As HX-Net uses a default stream velocity of 1 m/s, which would not be an accurate reflection of

the actual stream velocities, it is essential to compute manually the maximum allowable design

velocities. Correlations detailing these calculations are shown belowError! Bookmark not defined..

• Allowable velocity for given liquid =

( )2

1

liquidgivenofDensity

waterofDensitywaterforvelocityAllowable

where the allowable velocity of water in low carbon steel tubes is taken to be 10 ft/s.

Therefore, a sample calculation for HP steam generation is shown below:

Allowable velocity smsftsft /43.3/2.11lb/ft 49.4

lb/ft 62.4/10

2

1

3

3

==

=

• For gases and dry vapors in steel tubes,

Allowable velocity for gas (ft/s) = ( )( )

weightMolecularpsiainpressureAbsolute

1800

A sample calculation for Natural Gas feed stream is shown below:

Allowable velocity for NG feed = ( )( )

smsftpsia

/62.5/42.1845.162.580

1800==



The computed values for each process and utility stream were tabulated as follows in Table 1 and

Table 2. It can be seen that the liquid stream velocities were in the range of 1 to 4 m/s and vapor

streams were also well within the stipulated standard process design guidelines. Hence, we can

deduce that reasonable estimates have been obtained and these values were subsequently used in

place of the default value in HX-Net.

Table 1: Maximum design velocities for vapor streams.

Table 2: Maximum design velocities for liquid streams.

7.1.1.2 Determination of Flow Area Diameter

The flow area diameter in HX-Net was defaulted as 0.0254 m or equivalent to 1 inch. According

to heuristics1, 43 in. (19 mm) is a recommended trial diameter to commence design calculations.

Selection of a suitable tube thickness also depends on internal pressure and corrosion issues. For

plain carbon steels where severe corrosion is not expected, a minimum allowance of 2.0 mm

should be used.Error! Bookmark not defined. Hence a tube thickness of 14 BWG (wall thickness = 2.11

mm) was used. Therefore, considering the above factors, an internal flow area diameter of

0.01483 m was adopted for HX-Net computations.

Vapor Streams

Absolute Pressure

(psia)

Molecular Weight

Allowable Velocity

(ft/s)

Allowable Velocity

(m/s)

Pressure Category

Guideline VelocitiesError!

Bookmark not defined.

NG Feed 580.2 16.45 18.42 5.62 High

Atmospheric pressure:

10 to 30 m/s

High pressure: 5 to 10 m/s

SMR Feed 394.6 17.63 21.58 6.58 High SMR Outlet 381.5 12.67 25.89 7.89 High HTS Outlet 363.5 13.12 26.06 7.94 High LTS Outlet 360.5 13.12 26.17 7.98 High

Combustible Air

14.5 28.91 87.92 26.8 Atm

Flue Gas 11.5 27.07 102.02 31.1 Atm

Liquid Streams Density (lb/ft3)

Allowable Velocity

(ft/s)

Allowable Velocity (m/s)

Guideline VelocitiesError!

Bookmark not defined.

HP Steam Generation 49.4 11.2 3.43 1 to 4 m/s

Cooling Water 62.4 10.0 3.05



7.1.1.3 Calculations of Convective Heat Transfer Coefficients (HTC)

Of paramount importance is the verification of extracted HTC values against manually calculated

values computed from suitable correlations. This is because the magnitude of HTC values of

each component stream are used to compute the overall heat transfer coefficient, U, which is

then used to calculate the heat transfer area, an essential component of overall costing. Therefore,

it is essential to ensure the accuracy of the HTC values in order to obtain an accurate cost

estimation of the heat exchangers.

Heat transfer that occurs in a heat exchanger is classified under forced convection in a closed

conduit. For this type of flow, several correlations are applicable, dependent on whether the flow

is laminar or turbulent. After computing the Reynolds number for all the streams using the

maximum allowable velocity calculated in the previous section, it was found that the flow was

mostly in the turbulent region. For single phase turbulent flow either in gases or liquids, the

correlation proposed by Dittus and Boelter in 1930 can be readily applied2:

n

id

kh PrRe023.0 8.0=

where n = 0.30 if fluid is being cooled, and n = 0.40 if fluids is being heated.

To obtain the necessary dimensionless numbers in the above correlation, the arithmetic-mean

bulk temperature is used as the basis for evaluating stream properties. However, for the heat

exchanger involved in steam generation in the convection section of the furnace, there arises a

need to compute the convective HTC value for a 2-phase flow system where in-tube boiling

occurs. The Boyko-Kruzhilin equation3 was utilized to compute the HTC value as shown below

as it is generally conservative and adequately accurate for most purposes. Sample calculation for

steam generation in convection section of furnace:

( ) ( )

+=

2PrRe024.0 43.0

,

8.0

,

omim

ii

i

id

kh

ϑϑϑϑll

l

where ( ) i

v

v

im x

−+=

ϑϑϑ

ϑϑ l1 , ( ) o

v

v

om x

−+=

ϑϑϑ

ϑϑ l1



lk = thermal conductivity of the fluids, W/m °C; id = inner diameter of tube, m

lϑ = density of liquid phase, kg/m3, vϑ = density of vapor phase, kg/m3

ox = inlet vapor fraction, ix = outlet vapor fraction

( )( ) C.W/m294742

2.20.194.0240407

01483.0

58.0024.0 243.08.0 °=

+=ih

It was also noted that the heat exchanger cooling the LTS Outlet stream would result in the

condensation of that process stream into a 2-phase system too. The Boyko-Kruzhilin equation

was similarly applied to obtain its HTC value.

For the flue gas stream in the convection section, initial estimates of its Reynolds number seem

to indicate that the fluid is in transition flow. In this case, the HTC value would be bounded by

the laminar and turbulent conditions. Firstly, the HTC assuming a laminar flow regime was

calculated using the Hausen correlationError! Bookmark not defined.:

( )( )[ ]

14.0

32PrRe04.01

PrRe0668.065.3

+++=

wi

i

i

f

iLd

Ld

d

kh

µµ

where ih is the mean coefficient for the entire length L of a single tube, which was designated to

be 16 ft (4.88 m) long according to heuristicsError! Bookmark not defined. available in literature.

Neglecting the viscosity correction term

iw

i

,µµ

, ( )( )[ ]

+++=

32PrRe04.01

PrRe0668.065.3

Ld

Ld

d

kh

i

i

i

i

i

Another HTC value was then calculated as if the flow was turbulent using the Dittus-Boetler

correlation. The final estimated HTC value for transition flow would then be calculated as

follows: ( ) [ ]

−−+=

8000

2000ReiiiTi hhhh

7.1.1.4 Fouling Factors

When process and service fluids flow through a heat exchanger, fouling will usually occur, the

extent dependent on the nature of the fluid. The deposited material on the heat transfer surfaces

will normally have a low thermal conductivity and this will decrease the overall heat coefficient.

Hence the effect of fouling should be incorporated into the preliminary design, so as to oversize



the exchanger to allow for a performance reduction in actual operations. The list of fouling

factors used is included in the following

Table 3.

Table 3: Typical fouling factors for process and utility streams.

7.2 TARGETING

Besides achieving an integrated heat exchanger network which satisfies the heating and cooling

requirements of the various process streams, an important consideration is always an economic

one; the network design should be one that is a balance between cost and efficiency. Different

scenarios would warrant different approaches towards network synthesis. If utility costs are high,

a network which maximizes energy recovery within the plant would then be the optimal choice.

Conversely, if fuel costs are low, it would be an appropriate approach to opt for a network with

fewer heat exchangers so as to lower annual costs. Hence, an estimation of capital (purchased

equipment) cost and operating (utility) costs would have to be factored into the total annual cost

of the network, and subsequently, it would provide the direction and set the criteria for

evaluating each possible network design.

7.2.1 Cost Considerations

The total annual cost (TAC) of a network is calculated from the following formula:

CCAFOCTAC ×+=

where OC is the operating cost corresponding to the total utilities cost, AF is the annualization

factor and CC is the capital cost. The annualization factor is given by:

Type of Stream Typical Fouling Factors

(ft2.h.°F./Btu) Typical Fouling Factors

(m2.°C/W)

Natural Gas 0.001 0.000176

Light Hydrocarbon Vapors (clean) 0.001 0.000176

Hydrogen (saturated with H2O) 0.002 0.000352

Flue Gas 0.001-0.003. 0.002 was used 0.000352

Air (atmosphere) 0.0005-0.001. 0.00075 was used 0.000132

Steam (non-oil bearing) 0.0005 0.0000881

Cooling Water (treated makeup) 0.001 0.000176



( )

lifePlant

returnofRate

AF

+

=100

%1



The rate of return is 10% as stated in the problem statement. However, although the plant life

was stated to be 15 years, a plant life of 6 years was used instead after consultation with

Professor Rangaiah. This is due to the typical of the lifespan of heat exchangers. These values

were then keyed into HX-Net, where it will be used to automatically estimate the TAC with

appropriate OC and CC inputs.

7.2.2 Utility Cost Calculations

As the cost data given by the problem statement were in units incompatible with the input format

in HX-Net, there was a need to convert them to suitable units. Using high pressure (HP) steam

for sample calculations,

Cost of saturated HP steam (42 bar) = US$33 / metric ton = US$0.033 / kg

Temperature of Saturated HP Steam (42 bar) = 254.3°C

Mass heat of vaporization = 1726 kJ/kg

Plant operation time = 8000 h/year = 28800000 s/year

Cost index of HP Steam

=( )( )

onvaporizatiofheatmass

timeoperationPlantsteamHPofCost ( )( )kgkJ

yrskgUS

1726

28800000033.0$=

yrkWUS .6.550$=

Similarly, the cost of cooling water used was calculated as follows:

Cost of cooling water = US$0.067 / m3

Temperature range of cooling water = 90 to 120 °F = 32.22 to 48.89 °C

At 32.22 °C, Mass heat capacity = 4.314 kJ/kg °C, Mass density = 1002 kg/m3

At 48.89 °C, Mass heat capacity = 4.320 kJ/kg °C, Mass density = 989.1 kg/m3

Taking the arithmetic mean,

Mass heat capacity = 4.317 kJ/kg °C, Mass density = 995.55 kg/m3

Cost index of cooling water

= ( )( )

( )( )( )differenceeTemperaturcapacityheatmassdensityMass

timeoperationPlantwatercoolingofCost

( )( )( )( )( )CCCkgkJmkg

yrsmUSooo 22.3289.48317.455.995

28800000067.0$3

3

−−= yrkWUS .9.26$=



7.2.3 Heat Exchanger Capital Cost Estimations

Capital cost estimations have to be inputted into HX-Net so as to facilitate cost target

calculations. The capital cost index is represented in HX-Net as: cbAaCC +=

where CC = capital cost and a, b and c are constants.

Based on 1986 costs, the capital cost coefficients were given as a = 30800, b = 750 and c =

0.81.4 Any cost data must take into account the effect of time on purchased equipment cost, and

this was carried out by using the Chemical Engineering Plant Cost Index (CEPCI) to incorporate

inflationary effects over time5.

CEPCI for 1986Error! Bookmark not defined. = 318, CEPCI for Nov 20076 = 526

Therefore, correcting for changing economic conditions over time:

( ) 81.081.0 12415094675030800318

526AACC +=+

=

7.2.4 Supertargeting

Generally, as the minimum approach temperature, ∆Tmin, increases, a dual effect is seen. Firstly,

the amount of heat recovered from hot process streams decreases, leading to an increased need

for utilities, thus raising associated operating costs. However, having an increased ∆Tmin would

also mean more effective heat transfer, as the temperature gradient, acting as the driving force,

would be greater. Therefore, a smaller heat transfer area for the same target load would be

achieved, giving cost savings in terms of reduced capital investment in smaller heat exchangers.

Figure 1: (Left) Total cost and operating cost as a function of ∆Tmin. Figure 2: (Right) TAC and OC as a function of ∆Tmin.



Supertargeting using the cost indices that were computed earlier (capital and operating costs)

was carried out. It allows us to select the optimal ∆Tmin corresponding to the minimum TAC for a

network. As seen from Figure 1, the optimal ∆Tmin is in the region of 4°C. However, this was on

the low side and could be attributed to the faster rate at which the operating cost is increasing as

compared to the savings in the capital cost as ∆Tmin increases. Hence, to provide a sufficient

driving force, a ∆Tmin value of 10°C was chosen instead, since this would conform more to

industrial standards and also lead to only a marginal increase of TAC (less than 1%).

7.2.5 Comparison between the usage of HP and LP Steam Generation

It was found that TAC was actually a negative quantity, indicating that the designed network was

generating profits instead, due to the sale of high-value HP steam. Hence, it was decided that a

choice should be made to either generate HP or LP steam. After making the necessary

computations and keying the updated LP steam properties into HX-Net, a range target graph was

generated, as seen from Figure 2. It can be seen that the TAC has dropped significantly as

compared to Figure 1 where HP steam was used. Furthermore, ∆Tmin decreased to an even lower

value. Hence, it was decided that the usage of HP steam would be a better economical choice.

7.2.6 Calculation of Utility Targets

Once the heating and cooling requirements of the process streams have been determined together

with a specified ∆Tmin, the minimum amount of utilities can then be estimated. The temperature-

interval analysis, also known as the problem table algorithm, can be used for such a purpose.Error!

Bookmark not defined. Alternatively, the composite curve method first proposed by Umeda et al.7, can

be applied.

Figure 3: (Left) Hot and cold composite curves.

Utility Pinch



Figure 4: (Right) Utility composite curve showing the utility pinch created by HP steam generation.



As seen from

Figure 3, the graph is reminiscent of a threshold problem which only requires a cold utility,

unaffected by ∆Tmin. Figure 4 shows the utility composite curve where HP steam generation

and cooling water were used as the cold utilities. It is evident that the usage of HP steam

generation actually introduces a utility pinch at around 253.3 to 254.3°C, the temperature of

saturated HP steam at 42 bar. Hence the initial threshold problem with no pinch has now evolved

into one with a utility pinch, where pinch analysis can be applied by treating the utility streams

as “dummy” process streams8.

Figure 5: Network targets for ∆Tmin = 10°C.

7.3 MER NETWORK DESIGN

Although there are no process pinches in this design problem, a utility pinch has been introduced

with the usage of HP steam generation. A method, first introduced by Linnhoff and Hindmarsh9,

would be used to design two networks of heat exchangers, one on the hot side (above pinch) and

one on the cold side of the pinch (below pinch). Certain rules would have to be followed

according to the abovementioned method1:

1. The heat exchanger network is designed from the most constrained point where the approach

temperature difference is at its minimum, i.e. at ∆Tmin.

2. At the pinch, streams are paired such that above the pinch, Cp, cold ≥ Cp, hot and below the

pinch, Cp, hot ≥ Cp, cold.

3. The heat duty of each interior heat exchanger is selected to be as large as possible, so as to

reduce the total number of exchangers.



4. Hot utilities are added up to meet the heating energy targets, and no cold utilities are used

above pinch.



5. Cold utilities are added up to meet the cooling energy targets, and no hot utilities are used

above pinch.

The conditions for feasible pinch matches are as follows:

Number of streams going out of pinch ≥ Number of streams going into pinch. This is to

ensure that there are adequate streams for stream matching as certain utilities are not

permitted above and below the pinch.

MCp of the out streams ≥ MCp of the in stream. This is to ensure no minimum driving

force violations.

7.3.1 Stream matching above pinch

The number criterion is fulfilled here as there are 2 out streams (including the HP steam

generation utility) and 2 in streams. Although the only process cold stream (SMR Feed, circled

red in Figure 6) above the pinch should have been matched with the Flue Gas stream according

to the MCp criterion, this match was purposely left out as it was decided that the all the heat

energy above the pinch should be used for HP steam generation in the convection section. This

would not only give rise to more profits obtained from exporting the HP steam credit, but also,

the steam generation tubes in the convection section can act as shield tubes in the first few rows

of the convection section. This is because these tubes are exposed directly to radiation from the

radiant section. During actual furnace operations, there will be fluctuations in the temperature of

the flue gas; hence these tubes can thus act as a heat sink to regulate the temperature of the

exiting flue gas.

Excluding the above exception, the rest of the streams were matched so long as the head loads

were comparable and no driving force violations exist. MCp criterion need not may obeyed in

this non-pinch matches.

The pinch point has been calculated by HX-Net to be 263.3 / 253.3°C.

7.3.2 Stream matching below pinch

The number criterion is also fulfilled here with 2 out streams versus 1 in stream. A pinch match

was made between the HTS outlet stream and the SMR feed stream (circled yellow in Figure 6),



following the MCp criterion. Similar to the above pinch scenario, the rest of the matches were

then matched accordingly using similar heat loads and no driving force violations as the basis.

Stream splitting was only carried out for the HP steam generation stream. However, it was

represented in this way, rather than as separate utility streams, for network simplicity and easy

summing up of the total cooling load needed above the pinch.

7.3.3 Number of units in MER network

After stream matching has been completed for both sides of the pinch, the total number of heat

exchangers in the MER network was compared with the calculated minimum number of units

using the following equation:

1min −+= up nnu

where pn represents the number of process streams and un the number of utilities.

Minimum number of units above utility pinch 4114 =−+=

Minimum number of units below utility pinch 6116 =−+=

Total number of heat exchangers in the HEN network 1046 =+=

Therefore, MER condition has been met in the designed heat exchanger network as seen from

Figure 6.

7.3.4 Alternative MER Network Designs for Consideration

Two other MER networks were also designed to see if alternative MER networks would give a

lower TAC.

7.3.4.1 Network 1a

As seen from Figure 7, the heat exchanger network above the pinch remained the same. However,

a variation below the pinch was adopted. The combustible air to the furnace was solely preheated

by the convection section, instead of being heated by the LTS outlet stream, HTS outlet stream

and the convection section. This was because of considerations over the physical limitations to

the placement of the heat exchangers as increased pipe lengths and insulation would then have to

be factored into the cost of the network. However, it can be seen from the network performance

and cost indices table in Figure 7 that this network would not offer a lower TAC. Instead, capital

cost and total area would be higher than that of the original network design (Network 1). Since it

cannot be ascertained at this preliminary design stage that piping cost would be a major factor in



TAC and the units that make up a hydrogen plant are usually located close by, it has been

advised that Network 1 would suffice at this juncture.

7.3.4.2 Network 1b

As seen from Figure 8, again the heat exchanger network above the pinch remained in a similar

configuration as the previous 2 networks discussed. The combustible air was now preheated

solely by the LTS outlet stream. The TAC and total area calculated by HX-Net seemed to

indicate that this network was at a similar performance level as that of Network 1.

However it was apparent that cooling water would have to be used to cool the hot flue gases,

which would not be feasible as the primary use of the convection section should be to heat up the

other process cold streams. Cooling water should only be used as a last resort unless stack gas

temperatures are above environmental regulations. Hence this alternative was also discarded.

The network performance and cost indices of each MER network are summarized in Table 4.

The total areas for the 3 networks were satisfactorily within 20% of the target area. Network 1 is

thus selected for network evolution.

Table 4: Network performance and cost indices of each MER network.

Network 1 Network 1a Network 1b

HEN % of

Target HEN

% of Target

HEN % of

Target

Capital (US$) 7.540e+6 114.1 7.717e+6 116.8 7.533e+6 114.0

Total Cost (US$)

-4.819e+7 100.6 -4.813e+7 100.7 -4.819e+7 100.6

Total Area (m2) 1.774e+4 82.67 1.836e+4 85.56 1.769e+4 82.41



Figure 6: Network 1, original MER network design



Figure 7: Network 1a, air preheated solely by furnace convection section



Figure 8: Network 1b, air preheated solely by LTS outlet stream



7.4 NETWORK EVOLUTION

Evolution of the selected MER network (Network 1) was carried out, so as to reduce the

number of heat exchangers required, which would then result in a lowering of capital

costs. However, it is to be noted that the operating costs of the network would then

increase consequently as some energy recovery would have to be sacrificed and more

utilities would be needed. As a result, the TAC of an evolved network may be lower than

the MER network. Hence evolution of the MER network was done to facilitate a

comparison of the TAC of both networks.

7.4.1 Steps involved in network evolution

Determination of the scope of improvement.

Identification of loops present in exchanger network. An even number of units in

a loop have to be ensured

Breaking of loop, calculate resultant stream temperatures and check driving forces

Restore ∆Tmin if there is any violation by forcing heat along the path

Continue breaking other loops and restoring ∆Tmin one at a time if possible and

attractive

The number of minimum units assuming no pinch is 81271min =−+=−+= up nnu ,

which was found to be the same as the minimum number of units computed by HX-Net.

Hence the scope of improvement would be 2810 =− loops. These 2 loops were

identified by HX-Net and shown respectively in Figure 9 and Figure 11.

7.4.2 Evolution of 1st loop

The heat exchanger with the smallest heat load (circled green) was identified and its heat

load transferred to the next heat exchanger in the loop, as indicated by the green arrow in

Figure 9. The smallest heat load was eliminated because the cost savings in removing a

small exchanger would be much higher than the cost incurred in increasing the area of

another larger exchanger already in place. The temperatures of the affected streams were

then re-calculated with the new heat loads placed on them.



The evolved loop is shown in Figure 10. It can clearly be seen that there are now 2 heat

exchangers which have a temperature cross situation (circled white). Efforts were made

to restore ∆Tmin, however as there were no cold utilities already in place at the ends of the

2 process hot streams below the pinch, there was no viable way to prevent ∆Tmin from

being violated and a temperature cross situation to ensue. It would not be a good solution

to have to add a cooling water as a cold utility at the ends of either of the 2 streams, as it

defeats the purpose of evolving a network and reducing the number of exchangers. Hence

this loop was not considered for evolution.

7.4.3 Evolution of 2nd loop

Similarly, the heat exchanger with the smallest heat load (circled magenta) was identified

and its heat load transferred to the next heat exchanger in the loop, as indicated by the

magenta arrow in Figure 11. Again, as seen in Figure 12, a temperature cross was

encountered in one of the heat exchangers (circled white) due to the lack of cold utility

below the pinch, at the other end of the process stream. Hence, similar reasons prevented

the restoration of ∆Tmin, hence this loop was also not considered for evolution.



Figure 9: Network 1 showing 1st loop.



Figure 10: Evolution of 1st loop.



Figure 11: Network 1 showing 2nd loop



Figure 12: Evolution of 2nd loop.



7.5 HEAT EXCHANGER DESIGN

Shell and tube exchangers are the most commonly used type of heat transfer equipment in

the process industry, as they offer a large range of operating conditions, good mechanical

layout, large surface area in a compact setup, flexibility in choice of construction

materials, ease of fabrication and cleaning, and most importantly, well-established design

procedures. Standards of the American Tubular Heat Exchanger Manufacturers

Association (TEMA) are universally adapted. For petroleum and related industries with

generally severe duties, heat exchangers are typically fabricated in accordance with class

“R” TEMA specifications.10

7.5.1 Stream Data

With reference to Figure 6, the heat exchanger chosen for thermal sizing would be the

one using the SMR outlet stream to preheat the SMR feed stream (circled yellow). The

details of the unit are shown below:

Table 5: Heat exchanger stream properties.

Hot Stream Inlet Outlet Temperature 851.9 606

Effective Cp in kJ/kg°C 2.964

Thermal Conductivity, kf, in W/m°C 0.183

Density, ρ, in kg/m3 4.031

Viscosity, µ, in cP 0.02797

Mass flow rate, M, in kg/hr 183500

Fouling Factor in hr.sq. ft.°F /BTU 11 0.002

Fouling Factor in m°C/W 0.000352

Cold Stream Inlet Outlet Temperature 253.3 539.4

Effective Cp in kJ/kg°C 2.483

Thermal Conductivity, kf, in W/m°C 0.067

Density, ρ, in kg/m3 9.258

Viscosity, µ, in cP 0.02002

Mass flow rate, M, in kg/hr 183500

Fouling Factor in hr.sq. ft.°F /BTU Error! Bookmark

not defined. 0.001

Fouling Factor in m°C/W 0.000176



7.5.2 Material of Construction

The choice of a suitable material of construction is more often that not determined by

corrosion-related issues. Factors which influence the amount of corrosion include

oxidizing agents, solution pH, temperature, fluid velocity, films etc.12 For the heat

exchanger chosen for further thermal design, both the hot and cold streams are not of a

corrosive nature as hydrogen levels still has not made up a significant proportion of the

streams, and are also of low enough temperature such that temperature-related corrosion

is negligible. Even if fluid velocities are high, it can be overcome with design features

such as tube inserts or impingement baffles within the exchanger unit. After considering

the above factors, it was decided that carbon steel, the most common material for heat

exchangers, would be a suitable material for both the tube and shell sides.

7.5.3 Shell and Tube-Side Fluid Allocation

The following are considerations for fluid allocation in a heat exchanger with no phase

change10:

Table 6. Analysis for shell and tube-side fluid allocation

Factor General Guidelines SMR Outlet / SMR Feed Corrosion More corrosive fluid to tube-side to

reduce cost of expensive alloy or clad components

Both are non-corrosive

Fouling More fouling fluid to tube-side to increase fluid velocity, reduce fouling and facilitate cleaning

SMR outlet stream is more fouling due to the presence of H2

Fluid Temperatures Hotter fluid to tube-side to reduce heat loss, cost and for safety reasons

SMR outlet stream is hotter than SMR inlet

Operating Pressures Higher pressure stream to tube-side to reduce material cost

Both streams are of similar pressures

Pressure Drop Fluid with lower allowable pressure drop to tube-side to obtain higher heat-transfer coefficients

Both streams have similar allowable pressure drops

Viscosity More viscous fluid to shell-side, provided turbulent flow is achieved

SMR outlet stream slightly more viscous, but difference not significant

Stream Flow Rates Fluid with lower flow rate to shell-side

Both streams have similar flow rates



From Table 6, it was decided that the SMR outlet stream would be assigned to the tube-

side, while the SMR feed stream would be the shell-side fluid.

7.5.4 Exchanger Type

Designation of TEMA shell and tube heat exchangers follows a three-letter code to

specify the front end stationary head type, shell type and rear end head type. It was

decided that the AES type exchanger would be employed. Front head type A has a

removable channel and cover plate and is the least expensive option. Shell type E is the

most common shell construction but higher pressure drop may ensue. This can be

controlled by varying the tube layout and baffle pitch. Rear end head type S is most

commonly used for internal floating head designs, and allows for the cheapest straight

tube removal bundle.

7.5.5 Baffles

Baffles are used in the directing of fluid stream across the tubes so as to increase the fluid

velocity and thus improve the rate of transfer. For single segmental baffles, the maximal

baffle cut is about 45%. Optimum baffle cuts range from 20 to 25%, while the optimum

spacing is usually 0.3 to 0.5 times the shell diameter.10

7.5.6 Tube Dimensions

Typical tube diameters range from 85 in. (16 mm) to 2 in. (50 mm). A smaller diameter

( 85 to 1 in.) is usually used for most duties, as they give smaller and cheaper exchangers.

According to heuristics13, 43 in. (19 mm) is a recommended trial diameter to commence

design calculations. However, considering the large volume of cooling water needed, a 1

in. tube diameter (do = 25.4 mm, di = 21.18 mm) was employed.

Selection of a suitable tube thickness depends on internal pressure and corrosion issues.

For plain carbon steels where severe corrosion is not expected, a minimum allowance of

2.0 mm should be used.12 Hence a tube thickness of 14 BWG (wall thickness = 2.11 mm)

was used. Tube length L was designated to be 16 ft (4.88 m) long according to heuristics4

available in literature.



7.5.7 Tube Arrangements

Several tube arrangements exist: equilateral triangular, square or rotated square pattern.

An equilateral triangular pattern, as seen in Figure 13, was used as recommended by

heuristics, as it provides higher heat transfer rates. Furthermore, heavy fouling requiring

mechanical cleaning is not expected in this particular heat exchanger since shell-side

fluid is relatively clean. A tube pitch Pt (distance between tube centres) of 1.25do was

also utilized in accordance with common practices10.

Figure 13. Equilateral triangular tube arrangement.

7.5.8 Calculations

The log mean temperature difference lmT∆ is calculated as follows:

( ) ( )( )( )12

21

1221

lntT

tT

tTtTTlm

−−

−−−=∆ =

( ) ( )( )( )

Co2.332

3.2530.606

4.5399.851ln

3.2530.6064.5399.851=

−

−−−−

where lmT∆ = log mean temperature difference

T1 = hot fluid temperature, inlet

T2 = hot fluid temperature, outlet

t1 = cold fluid temperature, inlet

t2 = cold fluid temperature, outlet

To account for non-ideal counter-current flow within the exchanger, a correction factor Ft

is applied. The correction factor is a function of shell and tube fluid temperatures, and the

number of shell and tube passes, and is correlated as follows:

859.03.2534.539

0.6069.851

12

21 =−

−=

−

−=

tt

TTR 478.0

3.2539.859

3.2534.539

11

12 =−

−=

−

−=

tT

ttS

From the temperature correction factor plots for one shell pass, two or more tube passes10

or using a correction factor equation derived by Kern, the corresponding Ft = 0.883.



Since this value is above the benchmark value of 0.75 for an economic exchanger design,

it confirms that the single pass configuration is sufficient. The true temperature difference

mT∆ can then be computed:

CTFT lmtm

o3.2932.332883.0 =×=∆=∆

Since both hot and cold streams are gases, the overall heat transfer coefficient U should

be 10-50 W/m2°C. A value of 10 W/m2°C was used as the starting value for iteration.

The provisional required heat transfer area A would then be:

2

212300

3.293/10

10009.36158m

CCmW

kW

TU

QA

m

=×

×=

∆=

°°

Tubes with the following properties are chosen for use:

do = 25.4 mm

di = 21.18 mm

L = 4.88 m

As a first approximation, an allowance of 50 mm for tube-sheet thickness (2 tubes sheets)

was included, take mL 83.4= .

External surface area of one tube, ( )( ) 23 385.083.4104.25 mLdA ot =×== −ππ

Number of tubes needed, 3190031948385.0

12300≈===

t

tA

AN

Number of tubes per pass, assuming 2 passes, Np 15950231900 =÷=

Tube internal cross-sectional area, Ai ( ) 223 0003523.01018.214

m=×= −π

Area per pass = 22 62.50003523.015950 mmAN ip =×=×

Volumetric flow = density

hrperrateflowmass 1

3600×

smmkgs

hkg/65.12

/031.4

1

3600

/183500 3

3=×=

Tube-side velocity, smm

sm

passperarea

flowvolumetricut /25.2

62.5

/65.122

3

===



Check the tube-side velocity to ascertain reasonable value. The velocity is a bit too low

for high pressure vapor streams, where velocities should range from 5 to 10 m/s10. This

will be checked on later together with the pressure drop specifications.

7.5.8.1 Tube-Side Heat Transfer Coefficient Calculations

Reynolds number, Re 68701002797.0

02118.025.2031.43

=×

××==

−µρ it du

Prandtl number, Pr 453.0183.0

1002797.010964.2 33

=×××

==−

f

p

k

C µ

22818.21

1083.4 3

=×

=mm

mm

d

L

i

From Figure 12.23, the tube side heat transfer factor, 3101.4 −×=hj .

Tube side heat transfer coefficient then can be computed by 10:

=

−

i

f

w

hid

kjh

14.0

33.0PrReµµ

Neglecting the viscosity correction factor

wµµ

due to non-viscous nature of gases,

CmWhi

°− =

××××= 233.03 /18702118.0

183.0453.06870101.4

7.5.8.2 Shell-Side Heat Transfer Coefficient Calculations

For two tube passes in triangular pitch, the tube bundle diameter Db is estimated based on

the following empirical equation10:

mmmK

NdD

n

t

ob 235.55237249.0

319004.25

207.211

1

1

==

=

=

Since a split-ring floating head was used, the bundle diametrical clearance is

approximated as 78 mm10. Therefore, minimum shell inside diameter Ds = Db + 78 mm =

5315 mm = 5.315 m. This could be too large a value and would be corrected in the next

iteration by increasing the number of tube passes.

Kern’s “bulk-flow” method was used to estimate the shell-side heat transfer coefficient

and pressure drop. Although Kern’s method does not take into consideration the bypass



and leakage streams unlike the more vigorous Bell’s method, it was still adopted as it is

simple to apply and provides a satisfactory preliminary approximation.

An initial baffle spacing, lB equal to 0.4Ds, and a baffle cut of 25 percent was used to give

a balance of good heat transfer rates and minimal pressure drop.

Baffle spacing, mmmlB 126.2212653154.0 ==×=

tube pitch, pt mmm 03175.075.314.2525.1 ==×=

Cross-flow area, ( ) ( ) 26 23.210

75.31

212652374.2575.31m

p

lDdpA

t

Bsot

s =×××−

=−

= −

Mass velocity, ( ) 2/9.22

23.2

3600/1/183500mskg

shhkg

A

WG

s

s

s =×

==

Linear fluid velocity, smmkg

mskgGu s

s /47.2/258.9

/9.223

2

===ρ

The shell-side fluid velocity is lower than the recommended 5-10 m/s and would be taken

note in the next iteration.

Shell-side equivalent diameter, de (hydraulic diameter) for an equilateral triangular pitch

arrangement is computed as follows:

( ) ( )[ ] mmmmmmm

dpd

d ote 018.00.184.25917.075.314.25

10.1917.0

10.1 2222

0

==−=−=

Re 206001002002.0

100.189.223

3

=×

××==

−

−

µes dG

Pr 74.0067.0

1002002.010483.2 33

=×××

==−

f

p

k

C µ

Shell-side Nusselt number, Nu

14.0

31PrRe

=

w

Hj µµ

where jH is the dimensionless heat transfer factor.

For the calculated Reynolds number, the corresponding value of jH is 0.0042. 10

Neglecting the viscosity correction term

wµµ

,



Nu 3.7874.0206000042.0PrRe 3/13/1 =××== Hj

CmWm

CmW

d

kNuh

e

f

s

°°

=×

=×

= 22

/2910180.0

/067.03.78

7.5.8.3 Overall Heat Transfer Coefficient Calculations

The overall heat transfer coefficient is given by

iiidi

o

w

i

o

od hd

d

hd

d

k

d

dd

hhU

11

2

ln111 0

0

0

×+×+

++=

Where U = overall heat transfer coefficient, W/m2°C

ho = outside fluid film coefficient, W/m2°C

hi = inside fluid film coefficient, W/m2°C

hod = outside dirt coefficient (fouling factor), W/m2°C

hid = inside dirt coefficient, W/m2°C

kw = thermal conductivity of the tube wall material, W/m2°C

From literature, thermal conductivity of carbon steel = 55 W/m2°C 10

187

1

02118.0

0254.0

2839

1

02118.0

0254.0

552

02118.0

0254.0ln0254.0

5679

1

291

11×+×+

×

++=U

CmWUo2/3.95=

% error from estimated %853%10010

103.95=×

−=U

Since the overall heat transfer coefficient is not within 30% deviation from the initial

estimate, the design is not satisfactory and the calculated U would be used as the initial

value for the next iteration.10

7.5.8.4 Tube-Side Pressure Drop Calculations

Tube-side pressure drop is computed as:

25.28

2

t

m

wi

fpt

u

d

LjNP

ρµµ

+

=∆

−

Where jf is the dimensionless friction factor.



Ignoring the viscosity correction term,

( )[ ] 22

/2452

25.2031.45.22280052.082 mNPt =

×+×=∆

The calculated tube side pressure drop is below the specifications (below 3 psi =

20.68kPa, the designated pressure drop for heat exchangers with vapor service). The

number of tube passes should be increased in the next round of iteration.

7.5.8.5 Shell-Side Pressure Drop Calculations

Shell-side pressure drop is computed as:

14.02

28

−

=∆

w

s

Be

s

ft

u

l

L

d

DjP

µµρ

Ignoring the viscosity correction term,

22

/68202

47.2258.9

126.2

83.4

18

5315045.08 mN

m

m

mm

mmPt =

×

×=∆

Both pressure drops for the tube side and shell side are too low. The calculations would

have to be iterated again.

7.5.9 Modification of Design

As the overall heat transfer coefficient is not within 30% deviation from the initial

estimate, some modifications of the design parameters need to be implemented.

Firstly, 5 more iterations were done using the same overall parameters, each time using

the new calculated value of U. This generated a U that was within 30% of the previous

value. However, the value is 515.1 W/m2°C, still way much higher than the

recommended range of 10-50 W/m2°C. Furthermore, the shell side pressure drop was

more than the specifications.

Therefore, it was decided that a thicker tube with tube diameter (do = 50.8 mm, di = 46.59

mm) should be used. This yielded better results. After another 3 iterations, the value of U

decreased to 351 W/m2°C, however, the tube side and shell side velocities remained

above the 5-10 m/s guideline. A final change to the baffle cut from 25% to 45% was used.



After 2 iterations, it resulted in the following design specifications, which met most of the

specifications and remained fairly constant after a few iterations.

Table 7: Final heat exchanger specification.

This was taken to be the final heat exchanger specification as shown in Table 7.

7.5.10 Exchanger Cost

The cost of the single heat exchanger that was designed was estimated using the

CAPCOST program5. The bare module cost of a heat exchanger is computed as follows:

( )PMpBMpBM FFBBCFCC 21 +== °°

( )2

103102110 logloglog AKAKKC p ++=°

( )2

103102110 logloglog PCPCCFp ++=

Where Cp° = purchased cost, FM = material factor (carbon steel = 1), FP = pressure factor,

A = heat exchanger area (m2), P = design pressure (barg), B1, B2, K1, K2, K3, C1, C2, and

C3 are constants.

For a 1-2 shell and tube exchanger, made up of carbon steel tubes and shell, with a heat

exchanger area = 351 m2 and a maximum design pressure of 30 barg (an estimated 10%

safety factor was added on top of the normal operating pressure of 27 bar) , the bare

module cost as computed by CAPCOST is:

( ) ( )[ ]2

101010 351log3187.0351log8509.08306.4log +−=°pC

644,53$USC p =°

( ) ( )[ ]2

101010 30log0123.030log00627.000164.0log +−−=pF

04.1=pF

Tube Side Shell Side

Fluid velocity (m/s) 32.6 (5-10 m/s) 26.7 (5-10 m/s)

h (W/m2°C) 1650 576.3

∆P (kPa) 19.2 ( 20.68<∆P<258) 150 ( 20.68<∆P<258)

Overall

A(m2) 351

U (W/m2°C) 351

Q (MW) 36.1 (99.9% of target heat load)



( ) 488,176$1166.163.153644 USFCC BMpBM =××+== °

Given that CEPCI in 3Q 2001 = 39755 and CEPCI for Nov 20076 = 526, correcting for

inflation over year, taking into consideration that the above equations apply to the cost in

3rd quarter 2001,

000,234$835,233$488,176$397

526USUSUSCBM ≈=

=

7.6 RECENT DEVELOPMENTS

In recent years, there has been a growing interest in the research on the aspect of heat

exchanger fouling, with emphasis being placed on the methods of prevention and

cleaning. This is due to the high costs and time incurred with such operations. In order to

account for the effects of fouling in heat transfer, numerous models and simulations have

been developed to predict the rate of fouling but their accuracies are often limited by the

use of ideal fouling resistance that has few uses in real world applications. Several

measures have been adopted to combat fouling, an especially prevalent problem in

refineries as refineries seek to increase their profit margins, increasingly buying and

processing heavier, high sulphur and cheaper crudes, leading to elevated deposition

problems.

Strategies to ease problems includes13:

Feed analysis to minimize concurrence of two dominant fouling mechanisms

Blending light and denser crude oils together while avoiding precipitation

Further study of surface characteristics so as to fundamentally understand fouling

Use of intermittent pulsed flow

Feed at varying temperatures and pH values

These methods have shown promise in reducing fouling and had also resulted in

significantly improved HTCs that can allow small heat exchangers to be used instead.

Also, maintenance costs would be mitigated with these implementations.



7.7 HEAT EXCHANGER SPECIFICATION SHEET

Heat Exchanger data sheet Equipment number (Tag) A-123 Description/Function Hydrogen Plant

Type/Class AES / R No. of units 1 Shells per unit 1 Connected in (parallel or series) 1 1 Surface per unit (m2) 351 Surface per shell (m2) 351

Performance of one unit

SHELL SIDE TUBE SIDE Fluid circulating SMR Feed SMR Outlet Total fluid entering (kg/hr) 183500 183500

IN OUT IN OUT Vapor flow rate (kg/hr) 183500 183500 183500 183500 Liquid flow rate (kg/hr) - - - - Non-condensables flow rate (kg/hr) - - - - Temperature (K) 526 813 1125 879 Density (kg/m3) 9.258 4.031 Molecular weight 17.63 12.67 Viscosity liquid (kg/m.s) 2.002e-5 2.797e-5 Latent heat (kJ/kg) - - Specific heat (J/kg.K) 2483 2964 Thermal conductivity (W/m.K) 0.067 0.183 Operating pressure (kPa) 2700 2655 Velocity (m/s) 26.7 32.6 Number of passes 2 2 Fouling factor (m2.K /W) 0.000176 0.000352 Pressure drop (kPa) 150 19.2 Heat transferred (kJ/hr) 1.30e+8 MTD (corrected) (K) 293 Overall U (W/m2.K) 351

Construction of one Shell

Maximum operating pressure (kPa) 3000 3000 Maximum operating temperature (K) 813 1125 Type of unit Tube pitch 0.0635 m Joint Strength weld Tube material Carbon steel O.D (m) 0.0508 I.D (m) 0.04659 Length (m) 4.88 Shell material Carbon steel Diameter (Approx.) (m) 1.53 Tube Sheet material Carbon steel Baffle material Carbon steel Corrosion allowance (m) Tube side 0.00211 Shell side 0.00211 Baffle cross C.S. Type Segmental Spacing (m), % Cut 0.642, 45

Baffle arrangement

Nozzle arrangement

Remarks: Flared nozzle to reduce high inlet gas velocities.



7.8 INTEGRATED HEN WITH PFD OF PROPOSED HYDROGEN PLANT

Figure 14: Integrated heat exchanger network with PFD of proposed hydrogen plant.



APPENDIX A – STREAM DATA

Cold Stream Cold T In

(°C) Cold T

Out (°C) Hot Stream

Hot T In (°C)

Hot T Out (°C)

Load (kW) Area (m2) dT Min Hot (°C)

dT Min Cold (°C)

SMR Feed To Heated SMR Feed

253.3 539.4 SMR Outlet To Cooled HTS

Feed 851.9 606.0 36159 519 312.6 352.7

NG Feed To Heated NG

25.0 250.0 Flue Gas To Stack Gas 263.3 194.5 7308 3269 13.3 169.5

HP Steam Generation 253.3 254.3 HTS Outlet To Cooled LTS

Feed 418.3 263.3 23844 844 164.0 10.0

HP Steam Generation 253.3 254.3 Flue Gas To Stack Gas 565.6 263.3 32091 8859 311.3 10.0

Combustible Air To Preheated Air

25.0 80.3 Flue Gas To Stack Gas 194.5 150.0 4721 1368 114.2 125.0


218.0 253.3 HTS Outlet To Cooled LTS

Feed 263.3 234.3 4462 1554 10.0 16.3


80.3 106.1 HTS Outlet To Cooled LTS

Feed 234.3 220.0 2200 192 128.2 139.7

Cooling Water 32.2 48.9 LTS Outlet To Cooled PSA

Feed 217.7 50.0 29754 335 168.8 17.8

HP Steam Generation 253.3 254.3 SMR Outlet To Cooled HTS

Feed 606.0 353.9 37076 410 351.7 100.6


106.1 150.0 LTS Outlet To Cooled PSA

Feed 238.8 217.7 3747 393 88.8 111.6

Inlet Temp (°C)

Outlet Temp (°C)

MCp (MW/°C)

Enthalpy (kW)

HTC (W/m2°C )

Flowrate (kg/h)

Effective Cp (kJ/kg-°C)


Cold 218.0 539.4 0.1264 40621.1 479 183495 2.48

SMR Outlet To Cooled HTS Feed

Hot 851.9 353.9 0.1470 73234.72 471 183497 2.88

NG Feed To Heated NG Cold 25.0 250.0 0.0325 7307.54 764 42828 2.73

Flue Gas To Stack Gas Hot 565.6 150.0 0.1062 44118.78 42 312845 1.22


Cold 25.0 150.0 0.0853 10667.62 102 299742 1.02

HTS Outlet To Cooled LTS Feed

Hot 418.3 220.0 0.1539 30506.41 546 207402 2.67

LTS Outlet To Cooled PSA Feed

Hot 238.8 50.0 0.1774 33501.16 2074 207402 3.08



REFERENCE

1. Seider, W.D., Seader, J.D., Lewin, D.R. (2004). Product and Process Design Principles:

Synthesis, Analysis and Evaluation. 2nd Ed., John Wiley & Sons, Inc.

2. F.W. Dittus and L.M.K. Boelter, University of California, Publ. Eng., 2, 443 (1930

3. Bell, K.J., & Mueller, A.C. (2001). Wolverine Engineering Data Book II, Retrieved March

27, 2008, from Wolverine Tube Inc. Web site:

http://www.wlv.com/products/databook/databook.pdf

4. Hall, S.G., Ahmad, S., & Smith, R. (1990). Capital Cost Targets for Heat Exchanger

Networks Comprising Mixed Materials of Construction, Pressure Ratings and Exchanger

Types. Computers & Chemical Engineering, 14, 3, p. 319-335

5. Turton, R. et al (1998). Analysis, Synthesis and Design of Chemical Process. 2nd Ed.,

Upper Saddle River, NJ: Prentice Hall.

6. Economic Indicators. Chemical Engineering. Retrieved March 28, 2008 from Chemical

Engineering Web site:

http://www.che.com/business_and_economics/economic_indicators.html

7. Umeda, T., J. Itoh, J., & Shiroko, K. (1978). Heat Exchange System Synthesis. Chem. Eng.

Prog., 74, 70.

8. Shenoy, U.V. (1995). Heat Exchanger Network Synthesis : Process Optimization by

Energy and Resource Analysis. Houston: Gulf Pub

9. Linnhoff, B., & Hindmarsh E. (1983). The Pinch Design Method for Heat Exchanger

Networks. Chem. Eng. Sci., 38, 745

10. Chemical Engineering Design. 4th Ed., Oxford: Elsevier Butterworth-Heinemann

11. Branan, C. R. (2002). Rules of Thumb for Chemical Engineers : A Manual of Quick,

Accurate Solutions to Everyday Process Engineering Problems. 3rd Ed., Amsterdam; New

York: Gulf Professional Pub

12. Perry, R.H., & Green, D.W. (1997). Perry’s Chemical Engineers’ Handbook. 7th Ed.,

McGraw-Hill

13. Muller-Steinhagen, H. et al. (2007). Recent Advances in Heat Exchanger Fouling

Research, Mitigation, and Cleaning Techniques. Heat Transfer Engineering, 28, 3: pg. 173–

176

CN 4120: Design II Team 32 Economics & Profitability Report


Chapter 8 : COOLING TOWER


The cooling tower in this hydrogen plant is designed to provide a continuous flow of cooling

water required for the condensation and elimination of water vapour in the outlet stream of

low temperature shift (LTS) reactor, before it is fed into the pressure swing adsorption (PSA)

column for purification of hydrogen and removal of carbon dioxide. The cooling duty of the

tower is found to be 42.975 10 KW× . In order to meet this requirement, an induced draft

cooling tower with counter-flow pattern is designed.

The detailed design of this cooling tower would consist of the following sections:

• Physical dimensions of the cooling tower

• Cooling tower internals

• Material of construction

• Optimization and cost analysis



8.2 WORKING PRINCIPLES OF COOLING TOWER

Heat transfer in cooling towers occurs by two major mechanisms [1]: the transfer of sensible

heat from water to air by convection process and the transfer of latent heat by evaporation of

water.

Although sensible heat transfer due to temperature difference between the air and water

occurs, the extent of this heat transfer is much smaller as compared to the removal of heat

from water via latent heat of vaporization.(20% due to sensible heat:80% due to latent heat)

The governing equation for the heat transfer in cooling tower is the Merkel Equation [2],

defined to be

1

2

T

sa aTp

KaV dT

h hLC

−

− =−∫

where Ka = volumetric air mass transfer coefficient ( 3/ filllb air hr ft )

_

V = specific fill volume ( 3 2/fill Base Areaft ft )

_

L = loading factor ( 2

2 / Base Arealb H O ft )

hsa = enthalpy of saturated air at water temperature (Btu/lb dry air)

ha = enthalpy of air stream (Btu/lb dry air)

T1 = inlet water temperature ( oF )

T2 = outlet water temperature ( oF )

The derivation of this equation ignores the mass transfer resistance from bulk water to the

interface, the effect of evaporation, and the temperature differential between the bulk water

and interface. It demonstrates that the driving force for the cooling process is the enthalpy

potential difference between the interfacial film and surrounding air.[3]



8.3 Preliminary Design

Before we can proceed with the actual design of the cooling tower, the different

configurations for cooling tower must be examined for its advantages and disadvantages in

order to make a choice the configuration that is going to be used in this design report.

8.3.1 Selection of cooling tower

Cooling towers can be classified according to the means by which air is supplied to the

towers, i.e., natural draft vs. induced or forced mechanical draft (fans) and according to

relative movement of air and water, that is, counter-flow or cross flow.

8.3.1.1 Justification to reject the use of natural draft tower

Natural draft or hyperbolic cooling tower depends on the natural draft created by the

difference in the density of the entering and leaving air for movement. Due to their large sizes,

they are often used for water flow-rates above 200,000gal/min.[3] Though it does not incur

any operating or maintenance cost for fans and experiences almost no recirculation of hot air

that could affect tower performance, it is not used in this design due to the following reasons:

1. The construction of hyperbolic cooling tower requires large plot space which results

in higher capital investment on land

2. Natural draft tower depends completely on atmospheric conditions. This implies

that water temperature is difficult to control and maintained, which might affect

downstream units that utilize the cooling water. In our case, the affected unit will be

the heat exchanger that makes use of cooling water to eliminate water vapout [4]



8.3.1.2 Justification to use induced draft tower

Mechanical draft towers can be either forced or induced draft, depending on the position of

the fans. Forced draft towers have the fans located at the base of the tower which push the air

through the fill while induced draft units have fans located on the top of the tower which pull

the air through the packing and discharge them vertically upward at high velocity.

Induced draft cooling tower is preferred over forced draft towers due to following factors:[5]

1. Unlike forced draft tower which is subjected to recirculation of hot humid

discharged air into the fan intake, the recirculation issue is completely avoided in

induced draft since the air is discharged upwards at high velocity.

2. The induced draft allows a more uniform distribution of air inside the tower.

3. The power requirement of the fan system in induced draft tower is about half that of

forced draft tower for the same capacity.

4. Compared to forced draft tower, induced draft tower require less initial cost to start

up, take up less space and have the capability to cool over wide range.

8.3.2 Comparison between counter-flow and cross-flow Pattern

For cooling towers, counter-flow pattern is preferred over cross-flow pattern mainly because:

1. For a tower of similar capacity, 20-50% less pumping head is required for counter-

flow cooling tower as compared to cross-flow tower. This implies that counter-flow

tower can operate at a lower cost. [6]

2. Unlike cross-flow tower, counter-flow tower does not experience recirculation,

which greatly reduces tower performance due to higher wet-bulb temperature. [6]

3. Under the same design condition, a counter-flow tower produces more cooling per

unit volume at a lower cost.[6]

Based on the above factors, counter-flow pattern is chosen for my design of cooling tower.



8.4 DETAILED DESIGN OF COOLING TOWER

(All calculations in this section is based on Reference 2)

8.4.1 Specification of cooling tower design parameters

8.4.1.1 Wet bulb temperature

Wet bulb temperature is the temperature of the air entering the cooling tower and is the

lowest temperature at which water can be cooled to theoretically. Since a counter-flow

induced draft tower is designed, it is valid to assume that there is no recirculation. Hence, the

wet bulb temperature is taken to be that at ambient condition.

Taking a conservative approach, the ambient wet bulb temperature is determined using the

maximum dry bulb temperature in Singapore

As of 2007,[7]

Average daily maximum dry bulb temperature = 31.1oC

Mean relative humidity at 2 pm = 74%

From the psychrometric chart,

Ambient wet bulb temperature (t1) = 27.2 (81.0 )o oC F

8.4.1.2 Range

The range is defined to be the difference between the cooling tower water inlet (T1) and outlet

temperature (T2). Heuristic [8] assumes the maximum inlet temperature of cooling water to

be 120 (48.9 )o oF C and cooling water exit temperature to be90 (32.2 )o o

F C .

Hence,

Range = T1 – T2 = 120 – 90

= 30 oF = 16.7 o

C



8.4.1.3 Cooling water requirement

From the Heat Exchanger Network Design, the cooling water duty is found from simulation

using HX-Net to be 42.975 10 KW× . Hence, the amount of cooling water (L) required to

achieve this cooling load can be calculated by

1 2( )p

QL

C T T

=−

where Cp = Effective heat capacity of water at temperature 1 2

2

T T+.

42.975 10 3600

4.320(48.9 32.2)L

× ×=

− = 61.488 10 /kg hr×

8.4.1.4 Approach

Approach is defined to be the difference in temperature between the cooling water leaving the

tower and the ambient wet bulb temperature. As a general rule, the closer the approach to the

ambient wet bulb temperature, the more expensive the cooling tower due to increased size.

Approach = T2 – t1 = 90 – 81 = 9 oF = 5 o

C

This approach is very close to the typical approach of 10 15oF− in most cooling towers.

8.4.2 Exit air temperature and water to air flow ratio (L/G)

8.4.2.1 Exit air temperature

For a given set of cooling tower design conditions, an optimum design of the outlet air wet-

bulb temperature exists. This is desired as it will result in minimum construction and

operating costs. A good correlation exists between the optimum exiting air temperature (t2)

and the inlet and outlet cooling water temperature. This correlation which is to be used as a

rule of thumb for design is as followed:

1 22

120 90

2 2

T Tt

+ += = = 105 58.3o oF C=



8.4.2.2 Water to air flow (L/G) ratio

The L

Gratio of a cooling tower is the ratio between the cooling water and air mass flow rate.

Thermodynamically, the heat removed from the cooling water must be absorbed by the

surrounding air. Hence, the following energy balance can be used to evaluate the L

G ratio:

1 2 2 1( ) ( )pLC T T G h h− = −

2 1

1 2

( )

( )p

h hL

G C T T

−=

−

where h2 = Enthalpy of air-water mixture at the exit air temperature (kJ/kg dry air)

h1 = Enthalpy of air-water mixture at inlet air temperature (kJ/kg dry air)

Cp = Heat capacity of water at 25 oC

2 1

1 2

( ) (189.1 104.1)

( ) 4.18(48.9 32.2)p

h hL

G C T T

− −= =

− −

= 1.219

Assumptions made: 1) Cooling water mass flow is relatively constant (little evaporation)

2) Sensible heat transfer from water to air is negligible

8.4.3 Cooling tower characteristic

Tower characteristic can be determined by the Chebyshev method [1], whereby

1

2

1 2

1 2 3 4

1 1 1 1

4

T

sa aTp

T TKaV dT

h h h h h hLC

−

−

−= ≅ + + +

− ∆ ∆ ∆ ∆ ∫

where hsa = Enthalpy of air-water mixture at bulk water temperature (Btu/lb dry air)

ha = Enthalpy of air-water mixture at wet bulb temperature (Btu/lb dry air)

∆h1 = value of (hsa – ha) at T2 + 0.1(T1 – T2)

∆h2 = value of (hsa – ha) at T2 + 0.4(T1 – T2)

∆h3 = value of (hsa – ha) at T1 – 0.4(T1 – T2)

∆h4 = value of (hsa – ha) at T1 – 0.1(T1 – T2)



T, oF ( / )sah Btu lb dry air ( / )ah Btu lb dry air sa ah h− 1

h∆

90

93

102

108

117

120

55.93

60.28

75.42

87.76

110.60

119.54

44.80

48.46

59.43

66.74

77.71

81.37

∆h1 = 11.82

∆h2 = 15.99

∆h1 = 21.02

∆h1 =32.89

0.0846

0.0625

0.0476

0.0304

**Note that the enthalpy of air-water mixture increases 1 Btu multiplied by L

Gratio for every 1

oF of cooling.

( )120 900.0846 0.0625 0.0476 0.0304 1.688

4p

KaV

LC

−

−

−≅ + + + ≅

At this point, it should be noted that mechanical-draft cooling towers are usually designed for

L/G ratios ranging from 0.75 to 1.50 and the values of

p

KaV

LC

−

− vary from 0.50 to 2.50

consequently. Hence, the results obtained so far for the design of the cooling tower have been

satisfactory.

8.4.4 Loading factor

Loading factor (_

L ), also known as specific water flow rate, is the recommended cooling

water flow rate per unit volume of tower cross-sectional area. A general rule for loading

factor is that for difficult cooling jobs (large cooling range and/or close approach), a lower

loading factor is used and vice versa. A graphical method is presented to determine the

loading factor.



Method to obtain the loading factor from the sizing chart

1. A straight line is drawn first to connect the inlet and outlet water temperatures as

illustrated in the figure

2. Another line is drawn to intersect the first line at the wet-bulb temperature. This line.

would yield the water concentration or the loading factor L

** Note that the loading factor determined from this graph is lower than loading factor used with presently-used

fills. However, method for determining modern loading factor is proprietary information and is not available.

From the graph,

_

2 22.35 1176

min . .Base Area basearea

gal lbsL

ft hr ft= =

2

5746.

basearea

kg

hr m=

Assumption made: Density of cooling water does not change with temperature (1000 kg/m3)



8.4.5 Dimensions of Tower

8.4.5.1 Fill Height

The required fill or packing height Z is assumed to be equal to the specific fill volume _

V in

the Merkel equation. It can be calculated from the tower characteristic and loading factor.

The fill height (z) can be determined from the following equation:

( )calc

p

KaV Lz

KaLC

− −

−= ×

Past literature studies show that Ka value varied from 49 to 152 with 100 ± 30 as the average

value. Since Ka is proprietary information, Ka is assumed to be 100 based on previously-

designed cooling tower.

Hence,

1.688 117619.85 6.05

100z ft m

×= = =

8.4.5.2 Base area

The required base area or cross-sectional area (B) can be determined from the below equation:

621.488 10

2605746

LB m

L−

×= = =

8.4.5.3 Fill volume

The fill volume (V) can be calculated by

260 6.05V B z= × = × 31570m=



8.4.6 Make-up Water Requirement

In the cooling tower system, water makeup is required to replace the cooling water which is

lost through evaporation, drift and blow-down.

8.4.6.1 Evaporation loss (E)

The evaporation rate is assumed to be 1.0% of the water flow rate for each 10 oF temperature

drop through the tower. Hence, evaporation loss is calculated as shown below:

31 2 0.01 ( / )10

T TE L m hr

− = × ×

6

3

3

120 90 1.488 100.01 ( ) 44.64 /

10 10m hr

− × = × × =

8.4.6.2 Drift loss (D)

Drift is the entrained water in the tower discharge air. Drift loss is usually a function of the

drift eliminator design and typically varies from 0.1 to 0.2% of the water supplied to the

cooling water. Adopting a conservative approach, the drift loss is assumed to be 0.2% of the

circulating cooling water.

30.002 ( / )D L m hr= ×

6

3

3

1.488 100.002 ( ) 2.98 /

10m hr

×= × =

8.4.6.3 Blow-down (B)

Blow-down is the continuous or intermittent discharge of a small amount of the circulating

cooling water. Its purpose is to limit the increase in the solids concentration in the water due

to evaporation. Since chlorides remained soluble in the cooling water, the blow-down rate

can be determined from the cycles of concentration, which is the ratio of chloride content in

the circulating water to the chloride content in the makeup water as shown in the equation..

1( )

1B E D

Cycle= −

−



Cycles of concentration in cooling tower operations typically range from 3 to 5 but for

conservative design, it is assumed to be 3.

Hence,

3144.64( ) 2.98 19.34 /

3 1B m hr= − =

−

8.4.6.4 Makeup water requirement (M)

The makeup water requirement is the summation of evaporation loss, drift loss and blow-

down.

M E D B= + +

366.96 /m hr=

8.4.7 Power Requirement

Fans are essential for the function of induced-draft counter-flow cooling towers, so that air

can be forced to flow vertically upwards to be in contact and cool the process water. As this

circulating water is sprinkled down the tower by using nozzles, a pump is also required to

pump the water to the top of the tower for cooling. These two auxiliary units are the main

usage of energy for the operation of the cooling tower, and hence must be considered as a

factor in the design.

8.4.7.1 Pump power (Pp)

Pump power is determined from the following equation

61.98 10

p

p

L HP

η

×=

× ×

where L = Water flow rate (lbs H2O/hr

Hp = Pump head (ft)

η = Fan Efficiency (dimensionless, assumed to be 0.80)

10 19.85 10pH z= + = +

29.85 9.10ft m= =



Hence,

6

6

1.488 10 2.20462262 29.86

1.98 10 0.80pP

× × ×=

× ×

61.84 hp=

8.4.7.2 Fan Power (PF)

An estimate of fan power requirement is obtained from the volume of moist air moved by the

fan. For induced draft towers, this estimate is based on the exit air temperature which is

58.3 (105 )o oC F

At the air exit temperature,

Saturated absolute humidity of the air-water mixture:

H2 = 0.0507 lbs H2O/lb dry air

Density of dry air:

3

2

42.6439 42.64390.0755 /

460 105 460dry air lbs ft

tρ = = =

+ +

Density of water vapour:

( ) ( )3

2 2

26.6525 26.65250.9304 /

460 105 460 0.0507water lbs ft

t Hρ = = =

+ × + ×

Density of air-water mixture:

( )( )( )

( )( )( )

321 1 0.0507 0.0755 0.93040.0734 /

0.0755 0.9304

dry air water

mixture

dry air water

Hlbs ft

ρ ρρ

ρ ρ

+ × + ×= = =

++

Air flow rate:

1 6 -1 6( ) 1.488 10 2.20462262 (1.219) 2.69 10 /L

G L lbs hrG

−= = × × × = ×

Air flow rate (actual cubic feet of air per minute)

( )( )

( )( )

2 6 51 1 0.0507

(2.69 10 ) 6.418 1060 60 0.0734mixture

HF G acfm

ρ

+ += = × = ×

Assuming one hp is required for each 8000 actual cubic feet of air per minute moved by the

fan,

Fan power can be calculated by:

56.418 1080

8000 8000F

FP hp

×= = =



8.5 COOLING TOWER INTERNALS

8.5.1 Liquid Distributor

The liquid distributor serves the purpose of ensuring good liquid distribution by maintaining

uniform flow of liquid through the column. The type of liquid distributor used is very much

dependent on the flow pattern of the cooling tower.

A pressure type system of closed pipe and spray nozzles, like the one seen in figure 1, is

usually necessary for the counter-flow configuration.

Counter-flow distribution system in operation

Pressure spray system is more susceptible to clogging and more difficult to clean, maintain or

replace. However, it contributes significantly to overall heat transfer and does not require

high pump head in larger tower.

As for cross-flow tower configuration, the gravity flow distribution system is more

commonly used whereby the supply water is elevated to the hot distribution basin above the

fill. From this basin, the water flow over the fill (gravity-induced) through metering orifices

located in the distribution basin floor. Although this type of distribution system can be easily

maintained, it does not contribute to overall mass transfer and tends to require a higher

pumping head. There is also a tendency for algae growth if the basin is not covered.



8.5.2 Fill

The fill or packing is actually heat transfer surface that maximize water/air contact area and

increase the contact time between air and water for more effective heat transfer. Fill is

available in two types of design – the splash mode or the film mode.

Splash fill normally consists of horizontal slats in horizontal rows offset to one another to

cause the water to break up into droplets as it falls downward through the cooling tower. As

such, maximum exposure between the water surface and passing air is achieved. Splash fill is

characterized by reduced air pressure losses and is less conducive to clogging and can be

cleaned easily after a spill. However, it is very sensitive to inadequate support and must

remain horizontal and level. If sagging occurs, the water and air will channel through the fill

in separate flow paths, impairing the thermal efficiency greatly. [9]

On the other hand, film fill causes the water to flow in films over large vertical surface, thus

promoting maximum exposure to air. Film fill has the capability to provide more effective

cooling capacity within a given amount of space than splash fill. Film sheets are usually

spaced very close to each other. Due to the smaller passages, film fill is more sensitive to

plugging and makes the cleaning difficult if plugging do occurs. Hence in operations where

contamination by debris is possible, film fill should be avoided. [9]

8.5.3 Drift Eliminators

Drift Eliminators, as the name suggests, are used to remove entrained water droplets (also

known as drift) in the discharge air stream so as to prevent unnecessary water losses. The

separation is achieved by subjecting the discharge air to sudden change in flow direction.

Through the sudden directional change, a centrifugal force is created which cause the

entrained water droplet to deposit on the eliminator surface, from which it will flow back into

the tower.

Drift eliminator exists in many configurations but are usually classified according to the

number of directional changes or passes. More commonly found eliminators are the

“herringbone” type or the “honeycomb” type with labyrinth passage as shown in the figure

below.



“Honeycomb” type drift eliminator

Tighter control for drift release is expected nowadays due to possible environmental impact

associated with it. Drift can cause Legionnaires disease when mist droplets containing the

bacteria are inhaled into the body.

8.5.4 Supports

Although framework of cooling tower is already supported by massive cross-section, it is not

unbendable. Operations of large fans at high horsepower can result in large torsional forces

which could affect the stability of the tower. Therefore, it is necessary to construct a support

to maintain the proper positioning of the mechanical equipment used. These supports can be

in the form of heavy wall torque tubes welded to the outskirts of steel framework.

8.5.5 Cooling tower basin

The primary function of the cooling tower basin is to collect the cooled water leaving the

tower and to provide a reservoir for the cooling water pumps. In addition, it also serves as the

primary foundation for the tower and is also the collecting point for foreign materials washed

out of the air by water. Hence it must provide easy access for cleaning, have adequate

drainage facilities and be equipped with screening to prevent entry of debris into the suction

side piping.



To enhance the reliability of cooling tower, a minimum storage capacity should be provided

in the cooling tower basin to obtain the necessary time for corrective action during

emergency. Design practices worldwide recommend a minimum storage of 10 minutes.

8.6 MATERIAL OF CONSTRUCTION

A cooling tower must be able to withstand long duration dead loads imposed by the weight of

tower components, circulating water, snow and ice; as well as short term loads caused by

wind, maintenance and even seismic activities. Its design should be able to accommodate a

wide range of temperatures, a variety of external atmospheric conditions and internal pressure.

Corrosion caused by oxygenation and high humidity should also be taken into account.

Typical materials used are wood and steel. However, due to the above requirements of the

cooling tower, wood would not deem as a very suitable material. This is because although

wood is cheap and can last up to 30 years if it is well maintained, the drawback is that it is

susceptible to fungal and bacteria attack. Moreover, fungal and bacteria attack are more prone

to happen under the wet operating conditions of the cooling tower in a humid and wet tropical

country like Singapore. As a result of these factors, galvanized carbon steel was chosen as the

choice of material.

8.6.1 Liquid Distributor

Distribution systems are subjected to a combination of high temperature (hot water) and

oxygenation which are conditions favourable for corrosion. Hence, the material of

construction should be highly resistant to corrosion and erosion. Materials that are popularly

used include hot-dip galvanized steel, cast iron and redwood stave pipe. Because of the

relatively low pressure experienced by cooling tower piping, plastic can also be used for pipe

and nozzles construction. These plastic pipes are then reinforced with fibre for mechanical

strength.



8.6.2 Fills

Treated wood lath (primarily Douglas Fir) is considered as material of choice for splash type

fill due to its strength, durability, availability and relatively low cost. However, plastic such

as PVC which naturally has low flame spread rate is fast gaining popularity and dominance

due to safety consideration (fire-resistance properties). This is especially the case in steel

framed cooling water where fire-proofing is compulsory

Film fill, on the other hand, can be made of any material that is capable of being fabricated or

molded into shaped sheets, with a surface suitable for channeling of air and water. Currently,

the most popular material is PVC because of its chemical inertness, good strength and light

weight properties, low flame spread rate and most importantly, it can be molded to different

shape easily.

8.6.3 Drift eliminator

Similar to the fill, the material used for eliminator should be corrosion-resistant. In the

industry, treated wood and various plastic (predominantly PVC) are material acceptable for

drift eliminator.

8.6.4 Mechanical support

Traditional material used for these supports include carbon steel, hot dip galvanized after

fabrication or stainless steel at a significant additional cost. It is important to note that

stainless steel is not necessary as the combination of heavy construction and galvanization is

enough to meet the requirement for support.

Splash Type Fill – Wood Splash Bar Splash Type Fill – Plastic Splash Bars



8.7 COST ANALYSIS

8.7.1 Construction cost of for cooling tower

The number of tower units (TU) in a given cooling tower can be calculated by the following

equation:

( )TU L gpm Rating factor= ×

The rating factor is a measure of cooling job difficulty and a linear correlation exists between

the rating factor and the tower characteristic:

0.9964( ) 0.3843calc

p

KaVRating factor

LC

−

−= −

0.9964(1.688) 0.3843 1.298= − =

Hence,

61.488 10 4.40286754(1.298) 8504

1000TU

× ×= =

Cost of each tower unit is assumed to be US$14.45 (in 1978 dollars)

1978( $ ) 14.45 8504Construction Cost US = ×

$122,879.97US=

To correct 1978 dollars to 2007 dollars, we need to know the CECPI value for 1978 and 2007

CECPI (2007) = 528.2

CECPI (1978) = 218.8

2007 1978

(2007)( $ ) $

(1978)

CECPIConstruction Cost US US

CECPI= ×

528.2

$122,879.97 $296,641.68218.8

US US= × =



8.7.2 Operating Cost

The operating cost in the plant consists of two main components: the cost of makeup water

and the utility cost (electricity) that arises due to fans and pump operation.

8.7.2.1 Cost of makeup water

The cost of process water is assumed to be US$0.067 / m3

Plant operation time is taken to be 8000hours/year

Hence,

0.067 66.96 8000 $35,890.56 /Cost of makeup water US yr= × × =

8.7.2.2 Cost of Electricity

The total power requirement is the summation of pump power and fan power

62 81 143Total F pP P P hp= + = + =

Since electricity is sold at MW-h, the horsepower must be converted to MW-h

1 746hp W=

Hence,

6

143 746 3600143 384

10hp MWh

× ×= =

Since electricity cost is at US$100/MW-h

$38,400Electricity Cost US=



0

100000

200000

300000

400000

500000

600000

700000

95 100 105 110 115

Exit Air Temp (oF)

Co

ns

tru

cti

on

Co

st

0

10000

20000

30000

40000

50000

60000

95 100 105 110 115

Exit Air Temp(oF)

Op

era

tin

g C

os

t (U

S$

)

8.7.2 Optimization between the operating and construction cost

Optimization of the cooling tower is performed to minimize the operating and construction

cost of the cooling tower. The variable that is changed is the exit air temperature. From the

calculations shown above, it is clear that when exit temperature air temperature differ, the

L

Gratio, the tower characteristic

p

KaV

LC

−

− , the tower dimensions and the power consumption by

fan and pump will vary. The parameter that will not change during varying exit air

temperature is the makeup water requirement. Hence, it is not used as a guideline for

optimization.

Steps to perform optimization

1. An excel file is set up that contains all the equation used in the design of cooling

tower.

2. The exit air temperature is changed, and the construction cost and the operating cost

which is also the cost of electricity is monitored.

3. Five data point was tested, including the optimum temperature used as the initial

guess and the results were plotted in a graph

Chart of Operating Cost (US$) vs. Exit Air Temperature

Chart of Construction Cost (US$) vs. Exit Air Temperature



From these two plots, two observations can be made:

1. The operating cost, which is also the cost of electricity, decreases with increasing exit

temperature until around 43.3 oC (110oF) where it increases with increasing exit

temperature

2. The construction cost increases exponentially with increasing exit air temperature.

At the current moment, it is very difficult to consider whether 43.3 oC is the optimum

operating temperature because a lot of cost has not been factored into this optimization

investigation. These costs include the cost of auxiliary units such as fans and pumps which is

not included in the construction cost of the cooling tower, the cost for water treatment in term

of buying the additives.

With all these information, only then can we calculate the payback period to determine

whether it is worthwhile to switch to 43.3 oC. Hence, I shall adhere to the calculations in the

previous section.

8.8 ADDITIONAL CONSIDERATIONS TO COOLING TOWER DESIGN

8.8.1 Water Treatment

Cooling tower water treatment is necessary to minimize or eliminate: corrosion, scale and

biological fouling of heat transfer surface (in heat exchanger) which is caused by impurities

and minerals in the water. The difficulties caused by these impurities and minerals are

summarized in the table found in Appendix A [4].

8.8.1.1 Corrosion control

Corrosion is an electrochemical process that deteriorates metals exposed to water in the

presence of corrosive agents such as acids, oxygen, or bacteria. A common form of corrosion

is pitting.



Some of the possible causes of corrosion include:

1. Process leaks into the cooling water

2. water flow velocity that is too low (causes deposits and fouling lending to corrosion)

or too high (causes erosion and corrosion)

3. low pH and high temperature

To reduce corrosion to an acceptable level, chemical corrosion inhibitors which form

protective films on heat transfer surface are most effective. Examples of corrosion inhibitor

include phosphates, organics, zinc, nitrites, and molybdate salts. Unfortunately, the use of

chromate, which is a reliable corrosion inhibitor, is prohibited due to environmental

constraints.[4]

8.8.1.2 Scale control

Scaling is characterized by the formation of hard, dense deposits on material surfaces. These

deposits impact heat transfer. Calcium carbonate, which is formed from the reaction between

calcium ions and bicarbonate, is the main scaling constituent.

The key to prevention of scale formation in a cooling system is to maintain a reasonable

water velocity and to use chemical additives (dispersants) in combination with blow-down to

keep impurities concentration below the level which causes deposits. Equipment such as sand

pressure filter which requires minimal operating and maintenance cost can also be used.

8.8.1.3 Biological control

Operating conditions in cooling tower are ideal for growth of biological matters. These

conditions that encourage microbiological growth include favourable water temperature (20

to 50°C) and pH, continuous supply of nutrients and sunlight. If biological growth becomes

uncontrolled and form large sticky agglomerations, it may lead to operating problems as

listed below:

1. Fouling of heat transfer surfaces by bacterial slimes, resulting in flow restrictions

and high process temperatures.

2. Reduced cooling tower efficiency due to algae, fungi, and bacterial slime growth in

the water distribution basin and fill area of the cooling tower.



3. Corrosion

4. Clogging of water distribution nozzles.

Oxidizing biocides such as chlorine or sodium hypochlorite chlorine can be used to control

biological activity to prevent these operating problems from happening.

8.8.2 Environmental Concerns

Some of the environmental concerns with regards to cooling tower include:

1. Cooling tower blow-down is normally bypassed around major wastewater treatment

and discharged with treated wastewater.

2. Noise emission from fans and from the flow of cooling water over the tower may

require suppression if located near a community.

3. Spills and overflow of toxic and hazardous chemicals used for treatment of cooling

water must be contained.



8.9 CONCLUSION

In the sections above, the design of the cooling tower is carried out together with the cost

analysis and optimization. The specification obtained from the detailed calculations are

summarized in the table below

Cooling Tower Data Sheet

Tower Model Induced Draft

Flow Pattern Counter-Flow

Cooling Water Mass Flow Rate 1.488 x 106 kg/hr

Ambient Wet Bulb Temperature 27.2ºC

Exit Air Temperature 58.3ºC

On-Tower Cooling Water Temperature (Inlet) 48.9ºC (120 ºF)

Off-Tower Cooling Water Temperature (Outlet) 32.2ºC (90ºF)

Approach 5ºC

Range 16.7ºC

L

G Ratio 1.219

Tower Characteristic

p

KaV

LC

−

− 1.688

Loading Factor 5746 2.basearea

kg

hr m

Tower Dimension

Fill Height, z 6.05m

Base Area, B 260 m2

Fill Volume, V 1570 m3

Make-up Water Requirement

Evaporation (1% for every 10 ºF) 44.64 m3/hr

Drift Loss (0.02%) 2.98 m3/hr

Blow-down 19.34 m3/hr



Power Requirement

Pump Head 9.10 m

Pump Power 61.84 hp

Fan Power 80 hp Material of Construction

Cooling Tower Stainless Steel

Liquid Distributor Hot-Dip Galvanized Steel

Fill PVC

Drift Eliminator PVC

Mechanical Support Hot-Dip Galvanized Steel

Cooling Basin Concrete Cooling Tower Data Sheet



REFERENCES

1. Perry's Chemical Engineers' Handbook, Green, Don W. et al, McGraw-Hill, New York,

6th Edition, 1999, Chapter 12 2. Stephen A. Leeper, Wet Cooling Towers: ‘Rule-of-thumb’ Design and Simulation, U.S.

Department of Energy; Idaho National Engineering Laboratory, EGG-GTH-5775, July 1981.

3. Cooling Towers: Design and Operation Considerations, retrieved on 28 March 2008 from

Chemical Engineering Tools and Information website: http://www.cheresources.com/ctowerszz.shtml

4. G.B. Hill, E.J. Pring, Peter D. Osborn, Cooling Towers : Principles and Practice, London;

Boston: Butterworth-Heinemann, 3rd edition, 1990

5. J. D. Palmer, P.E., C.E.M. Evaporative Cooling Design Guidelines Manual. 2002 [cited 2008 March 10]; Available from: http://www.emnrd.state.nm.us/ECMD/Multimedia/documents/EvapCoolingDesignManual.pdf.

6. Donald R. Baker, Howard A. Shryock, A Comprehensive Approach to the Analysis of

Cooling Tower Performance, Technical Bulletin R-61 P-13, reprinted from the Journal of

Heat Transfer, August 1961. 7. Climate and Air Quality, in Yearbook of Statistics Singapore, National Environment

Agency. p. 4

8. Seider W.D., Seader J.D., Lewin D.R. Product & Process Design Principles. Edition 2

9. John C. Hensley, Cooling Tower Fundamentals, SPX Cooling Technologies, Inc.,

Overland Park, Kansas USA, 2nd edition, 2006, retrieved on 1 April 2008 from SPX Cooling Technologies website: http://spxcooling.com/en/library/detail/cooling-tower-fundamentals/



APPENDIX A IMPURITIES FOUND IN COOLING WATER

CONSTITUENT CHEMICAL

COMPOSITION DIFFICULTIES CAUSED

Hardness

+2Ca and +2Mg salts

expressed as CaCO3 Form scale deposit on heat transfer surface

Alkalinity Bicarbonate salts

expressed as CaCO3

Form calcium carbonate scales; attack

materials made of wood

Sulphate Sulphate ions −4SO

React with calcium in the water, forming

calcium sulphate deposits on condensers and

coolers

Chlorides Chloride ions −Cl

Add to dissolved solids content and increase

corrosion potential of cooling water

Silica Reactive Silica SiO2

React with calcium, magnesium and iron that

is in water to form silicate deposits

Ammonia Ammonium ion +4NH

Corrosion of copper and zinc alloys; Form

complex ions with zinc component in

corrosion inhibitors, rendering them

ineffective;

Dissolved solids ---

High concentration of dissolved solids causes

corrosion and increases precipitation of salts

which form scale deposits on heat transfer

surface

Suspended solids

(undissolved

matter)

---

Settling occurs when velocity decreases;

causes plugging, deposition in heat

exchangers and enhance biological growth

Oxygen and carbon

dioxide O2, CO2

General corrosion and local pitting of all

metal surfaces

Algae, bacteria,

fungi etc --- Organic growth and slime deposits



Chapter 9 : ECONOMICS & PROFITABILITY

9.1 INTRODUCTION

It is important to carry out a cost analysis on the design and selection of

equipment present in the chemical plant. Estimation of the initial cost of setting up the

plant is done following the development of the process using simulation software such as

HYSYS and MATLAB. Based on correlations and heuristics, we are then able to

determine dimensions of our equipment and consequently an estimated cost. In addition

to the cost of equipment, utilities cost can be obtained from HYSYS. A preliminary

gauge of the plant start-up cost can then be obtained. Following the finalization of the

plant design, we are able to approach a more realistic figure for the total cost, given the

availability of price information regarding the major units and auxiliary equipment from

the vendors.

An economic analysis follows next. Such an analysis is crucial because it enables

the management as well as the investors to assess the feasibility and profitability of the

plan before deciding whether or not construction of the plant should take place. Potential

returns of the plant are weighed against the risks that are involved.

For this part of the design project, total capital investments and total operating

cost are determined based on the correlation given by Turton [R1]. Revenues derived from

the sales of hydrogen are computed.

9.2 ASSUMPTIONS

1. The plant is to be located on Jurong Island, Singapore

2. The plant land is rented instead of being purchased. Rent rate (based on per

annum) is obtained from Jurong Town Council(JTC).

3. Operation time of the plant is 8,000 hours/year

4. Lifetime of the plant is 15 years, including construction time of 2 years.



5. Growth is at the same rate as the total cost of production

6. Salvage value for the plant is 10% of FCI.

7. All products of the steam methane reforming process are sold in the market.

8. The company is using its own profits to fund the investment cost of this project

and no bank loan is taken. It is assumed that the company has sufficient cash flow

to bear the investment cost and no interest would be paid to the bank. Hence

hydrogen that is produced can be sold at a more competitive price in the market.

In the event that the company wants to free up cash flow to sustain operations or

make other investments, a bank loan can be considered.

9. Corporate tax in Singapore is given as 18%.

9.3 CAPITAL COSTS

The total capital investment in the chemical process plant is made up of two main

components: the fixed capital and working capital [R1], i.e.

Total Capital Investment = Fixed Capital + Working Capital

The fixed capital represents all costs associated with the construction of the plant.

All fixed capital components are depreciable (except for land). The working capital

represents initial investment required to finance the initial phase of the operation before

revenues from the project starts. The working capital is usually used to pay wages, raw

materials and contingencies. As the working capital must be recovered at the end of the

project, it is a non-depreciable item on the

cash flow statement.

9.3.1 Computations for Fixed Capital

A list of the fixed capital costs is shown in Table 9-1. The fixed capital

investments include direct and indirect costs, costs for contingency and fee, as well as

auxiliary facilities costs.



Fixed Capital costs

Direct expenses

• Equipment purchase cost

• Materials used for installation cost

• Labour cost resulting from equipment installation

Indirect expenses

• Transportation, Insurance and taxes expenses

• Construction Overhead

• Contractor engineering expenses

Contingency and fee

• Contingency – for use in unpredictable circumstances

• Contractor fee

Auxiliary facilities

• Land Purchase

• Yard improvement

• Auxiliary development

• Offsite facilities and Utilities

Table 9-1: Items under Fixed Capital Costs

The direct and indirect expenses can be expressed in terms of the Bare Module

Cost (CBM). The CBM is the sum of all direct and indirect expenses incurred. To compute

CBM, the following equation is used.

BM

o

pBM FCC = (9-1)

where FBM is the Bare Module Cost Factor which accounts for the operating condition

and the material of construction. Cop is the purchased cost for base conditions (equipment

made of carbon steel and at atmospheric pressure)



C

op is given by Equation 9-2.

2

103102110 )(logloglog AKAKKCo

p ++= (9-2)

where A is the capacity parameter of the equipment.

When the capacity of the equipment lies outside the effective range of correlation,

the smallest possible capacity is used for cost calculations. For towers that have larger

volume than allowed, the costs are modeled as multiple columns in sequence.

To account for inflation, the fixed capital costs are inflated using the following formula:

bb I

I

C

C= (9-3)

where Cb is the known cost in the base year when the index was (=397 in 2001)

I (= 595.1 in December 2007) is taken to be the cost index in the year where the cost is C

To calculate the fixed capital costs, CAPCOST is employed. The costs for each

equipment is present in Table 9.2

CN 4120: Design II Team 32 Economics & Profitability


Table 9.2 Equipment Schedule

Section Tag No. in PFD

Equipment Name

Function Quantity Dimensions required for costing Cost, CBM

(USD)

LTS

R-103 LTS vessel

To further convert carbon monoxide and steam into hydrogen and carbon dioxide

1

Materials of construction Pressure Temperature Diameter Height

Low-Alloy Steel A387

1,370,000 25bar 220

oC

3.31m 5.07m

D-101 LTS knock-out drum

To remove condensate so as to prevent contamination of downstream catalyst

1 Materials of construction Low-Alloy Steel A387 542,000

Subtotal for LTS section 1,912,000

PSA V-101 PSA columns To purify hydrogen gas

8

Materials of construction Pressure Temperature Diameter Height

SS Clad

25,676,516 1bar-25bar 50

oC

3.00m 8.5m

Subtotal for PSA section 25,676,516

Cooling tower

V-102 Cooling tower To provide cooling water as a source of cold utility

1

Materials of construction Pressure Effective mass transfer area

Carbon steel

940,844 1bar

733m2

Subtotal for HTS section 940,844

Auxiliary units*

Heat exchangers

Expanders

4,895,966

4,476,000

Subtotal for Auxiliary units 9,371,966

CN 4120: Design II Team 32 Economics & Profitability


*Auxiliary units details are on following page: Heat exchangers Tag No. to PFD Cold streams Tube/shell

designation Hot streams Tube/shell

designation Area (m

2) Cost, CBM (USD)

HX-101 Expanded SMR Feed To Heated SMR Feed 1

Shell Cooled LTS Feed 1 To Cooled LTS Feed 2

Tube 1553.7 1,629,379

HX-102 Heated SMR Feed 1 To Heated SMR Feed 2

Shell SMR Outlet To Cooled HTS Feed 1

Tube 519.2 526,634

HX-103 HP Steam 1 Generation

Tube Cooled HTS Feed 1 to Cooled HTS Feed 2

Shell 409.7 581,212

HX-104 HP Steam 2 Generation

Tube HTS Outlet To Cooled LTS Feed 1

Shell 843.5 1,124,770

HX-105 Preheated Air 1 To Preheated Air 2

Shell Cooled LTS Feed 2 to Cooled LTS Feed 3

Tube 191.9 251,244

HX-106 Preheated Air 2 To Preheated Air 3

Shell LTS Outlet To Cooled Knockout Drum Feed 1

Tube 392.5 415,586

HX-107 Cooling Water Shell Cooled Knockout Drum Feed 1 to Cooled Knockout Drum Feed 2

Tube 335 367,140

Subtotal 4,895,966

Expanders

Item Item Quantity Material of construction Cost, CBM (USD)

E-101 SMR Feed Expander

1 SS 3,300,000

E-101 Steam

Expander 1 SS 360,000

E-105 H2

Expander 1 SS 816,000

Sub Total 4,476,000

Total Bare Module Cost = USD 78,680,946

CN 4120: Design II Team 32 Economic & Profitability


9.3.2 Computations for Total Module Costs

Besides accounting for the total bare module costs, it is also necessary to compute the

contingency and fee costs so as to account for the total module costs. The contingency and

contractor fee costs are assumed to be 15% and 3% of the bare module cost respectively. Adding

these two costs to the total bare module cost will give the total module cost. Alternatively, the

total module cost can be calculated from [R1]:

∑=

=n

i

iBmTM CC1

,18.1 (9-4)

The contingency and fee costs are tabulated in the following table.

Cost Item Cost(USD)

Contingency 11,893,090

Contractor Fee 2,378,618

Total Bare Module Cost 93,558,977

Total Module Cost 107,830,685

Table 9-10: Table for contingency and fee costs

9.3.3 Computations for Grassroots Costs (FCI)

The grassroots cost of the plant is calculated by adding the auxiliary facilities costs to the

previously calculated total module cost. The various auxiliary facilities costs is shown in Table

9-1. Since the plant designed is a new start-up, the grassroots cost is also equal to the fixed

capital investment (FCI).

As information on the various cost items is limited,grassroots cost is evaluated from [R1]:

∑=

+=n

i

o

TMGR iBMCCC

1,

50.0 (9-5)

= US$107,830,685 + 0.50 (79,287,268)

= US$147,474,319



9.3.4 Computations for Working Capital

The typical amount spent on working capital ranges between 15% and 20% of the fixed

capital investments, FCI. For conservative estimate, a value of 20% of fixed capital investment is

assumed for the working capital. Therefore

Working Capital = US$147,474,319 x 20%

= US$29,494,864

9.4 MANUFACTURING COSTS

To estimate the manufacturing costs involved in this chemical plant, there are 3

categories of costs that are included. They are as follows:

1. Direct manufacturing costs: These costs represent operating expenses that vary with

production rate. When product demand decreases, production rate is also dropped below

the design capacity and there would be a decrease in the factors making up the direct

manufacturing costs.

2. Fixed manufacturing costs: These costs are independent of changes in production rate.

They include property taxes, insurance and depreciation. These costs are charged at

constant rates even when the plant is not in operation.

3. General expenses: These costs represent an overhead burden that is necessary to carry

out business functions. These include management, sales, financing and research

functions.

4. Land lease cost: Since the land is rented, it will not be included in the fixed capital

investment, but in the operating cost of the plant.



Cost item Value

1. Direct manufacturing costs, DMC

Raw materials

Waste treatment

Utilities

Operating labor

Direct supervisory and clerical labour

Maintenance and repairs

Operating supplies

Laboratory charges

Patents and royalties

CRM

CWT

CUT

COL

0.18COL

0.06FCI

0.009FCI

0.15COL

0.03COM

Total DMC = CRM + CWT + CUT + 1.33COL + 0.03COM + 0.069FCI

2.Fixed Manufacturing Costs, FMC

Depreciation 0.1FCI

Local Taxes and Insurance 0.032FCI

Plant Overhead Costs 0.708COL + 0.036FCI

Total FMC = 0.708COL + 0.068FCI + Depreciation

3.General Expenses, GE

Administration Costs 0.177COL + 0.009FCI

Distribution and Selling Costs 0.11COM

Research and Development 0.05COM

Total GE = 0.177COL + 0.009FCI + 0.16COM

4. Land lease, CL

Total Costs = CRM + CWT + CUT + 2.215COL + 0.190COM

+ 0.146FCI + CL + depreciation

Table 9-10: Components for Costs Of Manufacture



9.4.1 Operating labour costs, COL

The number of operators per shift, NOL, is calculated using the following equation:

5.02 )23.07.3129.6( npOL NPN ++= (9-6)

where P is the number of processing steps involving the handling of particulate solids and Nnp

is the number of non-particulate processing steps handing steps and includes compression,

heating and cooling, mixing, and reaction. In general Nnp is given by:

∑= EquipmentNnp (9-7)

where equipment comprises of compressors, towers, reactors, heaters and exchangers, and

excludes pumps and vessels.

No. of towers 9

No. of reactors 3

No. of heaters 1

No. of exchangers 10

Nnp 23

Table 9-11: Number of equipment

Since the plant does not handle particulate solids, P = 0.

Therefore, NOL= (6.29 + 31.7(0)2 + 0.23(23))0.5 = 3.40

The following assumptions are made when calculating COL.

• The plant is operating 8000hrs/yr = 47.6weeks/yr

• Each operator works 5 shifts per week, and each shift is 8 hours, thus an operator

works 47.6 × 5 = 238 shifts per year

• Assuming plant operates 24hrs/day, there are 3 shifts in a day. Total number of

shifts per year = 8000/8 = 1000 shifts/yr

• Number of operators needed to provide this number of shifts = 1000/238 = 4.2

operators



Hence, operating labour needed = 4.2 × NOL = 14.3 ≈ 15

COL (2001) = 15 × US$50,000 = US$750,000 /yr

COL (2007) = US$1,004,691/yr

9.4.2 Utility costs, CUT

9.4.2.1 Electricity

Below is the calculation for the electricity cost for cooling tower.

Cooling tower Electricity (hp)

Pump power 173.95

Fan power 225.87

Power 399.82

Table 9-12: Total Power for Cooling Tower

Below is the calculation for the electricity cost for furnace.

Furnace Electricity (hp)

Induced draft fan 25

Forced draft fan 30

Power 55

Table 9-13: Total Power for Furnace

Total power (hp) Total power (MW) Cost of electricity Total annual cost

454.82 0.339 US$100/MWh US$ 271,200

Table 9-14: Total power for plant

Total electricity cost per annum (2007) = US$ 271,200



9.4.2.2 Cooling water cost

The cost for cooling water used in the cooling tower and the heat exchangers are

calculated as follows:

Water makeup Volume (m3/hr)

Evaporative Loss 124.99

Drift Loss 8.34

Blowdown 54.15

Total water makeup 187.49

Table 9-15: Process water used for cooling tower

Given cost of water is 0.067 US$/m3,

Cost of water used for cooling tower = $100,494

Total volume of water (m3/hr) Cost of water US$/m3 Total annual cost (US$/yr)

4280.71 0.067 2,258,570

Table 9-16: Process water used for heat exchangers

Volume of water used for heat exchangers = 4280.71m3/hr

Cost of cooling water used for heat exchanger (2007) = US$2,258,570

CUT = US$(271,200 + 2,258,570 + 100,500)

= US$2,630,270

9.4.2.3 Waste treatment costs, CWT

In this plant, the waste water coming out of the LTS knockout drum should be treated.

The volume of wastewater flowing out of this knockout drum is 92.23m3/hr. The cost of treating

this water is US$41/1000m3 of waste water [R1].

Cost of treating LTS knockout drum waste water (2001) = 92.23 × 8000 ÷ 1000 × 41

= US$30,250



CWT = US$(30,250 × (595.1 ÷ 394.3))

= US$45,657

9.4.3 Raw materials costs, CRM

The raw materials used in this plant are natural gas and high pressure steam.

Section of plant

Raw material

Consumption Cost Annual cost of material (2007)

Feed to plant Natural gas 2603 kmol/h US$7.59/kmol US$158,054,160

HTS reactor HP steam 23910 kg/h US$33/tonne US$6,312,240

To SMR HP steam 140700 kg/h US$33/tonne US$37,144,800

Table 9-17: Raw materials used for the plant

Unit Type of catalyst Cost of catalyst

Lifespan of catalyst

Annual cost of catalysts (2007)

HTS reactor Cr iron oxide catalyst

US$109414 3 years US$36470

LTS reactor Cu Zn oxide catalyst

US$226800 3 years US$75600

PSA reactor Activated carbon and zeolite 5A

catalyst

US$372372 3 years US$124124

Table 9-18 Cost for catalysts

CRM = US$(158,054,160 + 6,312,240 + 37,144,800 + 36,470 + 75,600 + 124,124)

= US$ 203,751,994

9.4.4 Land lease, CL

The land area required per year = 29400m2

From Jurong Town Corporation, the price of land rental at Jurong Island is $11.87psm per year.

CL = $29400 ×11.87 ÷1.3577

= US$257,036



9.4.5 Computation of manufacturing costs

To calculate COM, cost of manufacture, the following equations were used:

DMC = CRM + CWT + CUT +1.33COL + 0.069FCI + 0.03COM

FMC = 0.708COL + 0.068FCI + depreciation

GE = 0.177COL + 0.009FCI + 0.16COM

Thus, adding all 4 components together, we can solve for COM:

COM = 0.180FCI + 2.73COL + 1.23(CUT + CWT + CRM) + CL + depreciation

Cost item Cost (US$) 1. Direct manufacturing costs, DMC Raw materials

Waste treatment

Utilities

Operating labour

Direct supervisory and clerical labour

Maintenance and repairs

Operating supplies

Laboratory charges

Patents and royalties

203,751,994

45,657

2,630,264

1,004,692

180,845

8,848,459

1,327,269

150,704

8,938,259

Total DMC = US$226,559,293

2.Fixed Manufacturing Costs, FMC

Local Taxes and Insurance 4,719,178

Plant Overhead Costs 6,020,397

Total FMC = US$10,662,888

3.General Expenses, GE

Administration Costs 1,505,099

Distribution and Selling Costs 32,773,615

Research and Development 14,897,098

Total GE = US$49,071,189

4. Land lease, CL = US$257,036

Total Manufacturing Costs, COMd (without depreciation) = US$301,540,961

Table 9-19: Computed values for Costs Of Manufacture



9.4.6 Salvage value

Salvage value, S, represents the fixed capital investment of the plant minus the cost of the

land, at the end of the plant life. Assuming the salvage value of the property at the end of service

life is 10% of FCI.

Salvage Value, S = 10% of FCI = US$14,747,432

9.4.7 Depreciation

Depreciation is the reduction in value of equipment due to physical deterioration. Depreciation is

calculated using the straight line depreciation method as follows:

Depreciation , dk = [FCI-S]/n

= (147,474,319 - 14,747,432)/15

= US$8,848,459/yr

9.4.8 Revenues

Steam revenue

Steam is generated from the heat exchangers in the plant.

Steam generation revenue (2007) = US$67,987,243/yr

Electricity revenue

Below is the table showing the electricity generated by the expanders in this plant.

Assuming turbine efficiency = 70%,

Component Duty (kW) Electricity generated (kW) Revenue (US$)/yr

SMR feed expander 1787 1251 1,000,800

Steam expander 41.28 28.90 23,120

H2 expander 287.2 201.04 160,832

Total Revenue 1,184,752

Table 9-20: Total Electricity Generated Revenue

Electricity generated revenue (2007) = US$1,184,752/yr



Hydrogen revenue

Component Output (m3/yr) Output (kg/yr) Unit Price (US$/kg)

Hydrogen 1250,000,000 110,925,000 2.43

Total Revenue 270,024,351

Table 9-21: Total Hydrogen Generated Revenue

Total revenue from Hydrogen, Steam and Electricity

= US$ 67,987,243 + 1,184,752 + 270,024,351

= US$ 339,196,346

9.5 PROFITABILITY ANALYSIS

To find out the profitability and feasibility of a designed plant, an analysis of the cash

flow diagrams will be useful. Discrete and cumulative cash flow diagrams provide a clear insight

to the investments and profits which are made for every year of the plant project. The time value

of money is also important concept for assessing the profitability of a plant. The value of money

differs with time due to the earning capability of the money. The difference in the value of

money with time is not due to inflation and does not include the purchasing power of money.

This is an important concept as the designed plant is to operate over a span of several years and it

is hence more accurate for our profitability analysis if the time value of money is taken into

consideration.

In this section, the designed plant is considered to be built over 2 years and a plant life of

13 years follows, making up a total of 15 years for economic analysis. The cumulative cash flow

diagram for this project is studied. At the same time, the price at which the produced hydrogen is

to be sold so as to obtain a 10% return on the investment on the plant over a total payback period

of 15 years will be investigated.



9.5.1 Land Cost

The land cost is usually an investment that is made to the plant at the building stage of the

plant. The piece of land may be rented for the purpose of building and running the plant over its

construction + operation lifespan, which is taken to be 15 years for this project. In this case, the

land is being rented for the plant to be located in Jurong Island, Singapore. Jurong Town

Corporation, JTC has provided the rental cost which is US$8.74/m2 for each rental year.

Taking account of the dimensions of the various units as well as the amount of land

required for auxiliary facilities, the total required land area of 29400m2 for our plant will cost

US$257,036/yr. This amount of yearly rental is to paid from the construction of the plant till the

end of the plant operating life.

9.5.2 After Tax Cash Flow

The sales revenue made from the produced hydrogen, electricity and steam is not the

profit made to the plant due to expenses such as manufacturing costs and depreciation as well as

payable income tax. The net cash inflow to the plant or cash profit to the plant is thus after tax

cash flow here the cost of manufacture, cost of land rental, depreciation and income tax have

been taken into account.

The price of the hydrogen would be determined in the profitability analysis to assure a

rate of return of 10% to the plant. Since the plant is to be built and operated in Singapore, the

taxes payable would follow the corporate tax regulations on profits in Singapore by the Inland

Revenue Authority of Singapore, IRAS. The tax regulations state that corporate tax rate would

be at 18% from 2008 onwards and tax exemption is only valid for companies with income lesser

or equal to S$300,000 [2]. Since the positive income generated by the plant yearly is more than

S$300,000, a tax rate of 18% will be imposed on the designed plant.

t=18% (9-8)



This tax scheme is used to calculate the amount of tax payable. As income refers to the

net income to the organization, the percentage tax should be multiplied to the net value of

revenue taking away expenses which are the cost of manufacture without depreciation.

Income tax = (R - COMd –d)t (9-9)

The after tax cash flow to a plant is therefore

After tax cash flow = (R - COMd –d) (1- t) +d (9-10)

9.5.2 Rate of Return on Investment (ROROI)

The rate of return on investment of a project can be determined from the ratio of the

average net profit to the fixed capital investment excluding land cost. The average net profit can

be obtained from averaging the cumulative cash flow at the end of the project (after 15 years).

As the ROROI refers to the rate of return on investment in a project, a positive ROROI is

expected for a feasible project, hence the ROROI should ideally be as high as possible. ROROI

has been set to be 10% for this project.

9.5.3 Net Present Value (NPV)

The net present value is the cumulative discounted cash flow at the end of the project. A

non-negative NPV is required for a feasible project and hence is assumed to be zero. A non-

negative NPV indicates that the plant would at least have a rate of return that is equal to the

discount rate used in the calculation. For this work, we employ a NPV of 0 to back-calculate the

price of hydrogen which we should be selling at in the market.

9.5.4 Discounted Cash Flows in Project

Keeping the above cash flows of the project in mind, a cash flow table of the project can

be determined once the selling price of the hydrogen is set. To ensure a 10% return on the



investment for the discounted cash flow for a payback period of 15 years, a price of the hydrogen

is set such that the return on investment for the discounted cash flow is 10% and taking the

cumulative discounted cash flow at the end of the whole project to be zero. The ‘Goal-Seek’

function under Microsoft Excel is employed for this purpose. Table 9-22 shows the cumulative

discounted cash flows for our designed plant.

End of year, k

Investment Depreciation Gross Profit

After tax cash cash flow

Discounted Cash Flow

Cumulative Discounted Cash

Flow

0 -260599 - 0 -260599 -260599 -260599

1 -88484591 - 0 -88484591 -80440538 -80701137

2 -88484591 - 0 -88484591 -73127761 -153828898

3 0 8848459 28209037 24724133 18575607 -135253291

4 0 8848459 28209037 24724133 16886916 -118366375

5 0 8848459 28209037 24724133 15351741 -103014634

6 0 8848459 28209037 24724133 13956129 -89058505

7 0 8848459 28209037 24724133 12687390 -76371116

8 0 8848459 28209037 24724133 11533991 -64837125

9 0 8848459 28209037 24724133 10485446 -54351679

10 0 8848459 28209037 24724133 9532224 -44819455

11 0 8848459 28209037 24724133 8665658 -36153797

12 0 8848459 28209037 24724133 7877871 -28275927

13 0 8848459 28209037 24724133 7161701 -21114226

14 0 8848459 28209037 24724133 6510637 -14603589

15 44242296 8848459 72451333 61002816 14603589 0

Table 9-22: Cummulative discounted cash flows

By fixing cumulative discounted cash flow to be 0 at the end of 15 years, a selling price

of US$2.43/kg is obtained via the ‘Goal-Seek’ function.



Fig 9-1: Cumulative Discounted Cash Flow Plot

As we can see from the graph above, a selling price of $2.43/kg for the hydrogen product

will enable us to get a cumulative discounted cash flow of zero after 15 years (2 Construction

years + 13 Operation years).

We next study the price of hydrogen which we have to sell at, if we are concerned with a

shorter discounted payback period, as illustrated in Fig 9-2.

As we can see from figure 9.2, if we would like to recover our capital investment just

after 1 year of operation, we will need to sell hydrogen at US$4.19/kg. An important assumption

made here is that all hydrogen that is produced will be sold and there will be no leftovers. Given

more years for the returns, the price of hydrogen can be set at a lower rate. This is better overall

so that hydrogen can be priced more competitively, in view of numerous other hydrogen

providers in the market.



Fig 9-2: Price of Hydrogen over the number of years of operation

9.6 FEASIBILITY OF STORAGE FACILITIES FOR NATURAL GAS FEED

This portion of the report addresses the feasibility of an installation of storage tanks for

the natural gas feed, as a recommendation for future consideration.

For our design at present, the natural gas feed has been assumed to be provided by

vendors via pipes and on top of this, there will be no interruption in the provision of natural gas

by the vendors. Product hydrogen is not stored and exported immediately once produced, since

we have earlier made the assumption that every unit quantity of hydrogen will be able to find its

customer in the market.



Realistically speaking however, it is unlikely that there will be perfect provision of

natural gas without any failures everyday around the year. 2 possible scenarios of failures in the

provision of natural gas are identified:

1. Temporary complete-termination of natural gas provision due to errors at vendor’s end

2. Fluctuations (less/more than the contracted amount) in the natural gas provision

Both scenarios result in potential losses for the plant. Hence installation of storage tanks for

stand-by natural gas feed can be a solution around these problems.

In making a cost analysis of the proposed storage tanks, the following assumptions are

made:

1. Interruptions in natural gas provision do not last not more than 3 days, hence designs for

3 days’ worth of natural gas feed are made

2. Should there be any identified emergency in natural gas provisions, stand-by natural gas

from the storage tanks will be utilized with immediate effect and there is negligible delay

associated with the operation of control components

3. Heaters, coolers and piping constitute 20% of tank equipment cost

4. Tanks are of floating-roof type

Given that 2603 kgmole/h gaseous natural gas feed is provided by the vendor at 25oC &

40 bar, the following has to take place in order to store liquid natural gas in the storage tanks:

Fig 9-3: Storage process of natural gas feed



1. Natural gas feed at 25oC will be cooled to -87oC to transform the gaseous feed into liquid

form to facilitate storage in the tank. Temperature at which the gaseous feed will

liquidize, i.e. -87oC has been verified with HYSYS.

2. In the event of an emergency such that the stand-by natural gas has to be employed, the

liquidized form of natural gas will be heated up slowly and gradually, to avoid the

potential hazard of rapid expansion due to vaporization, and rupturing of the pipes as a

result. Natural gas in gaseous form can then be fed into the furnace.

The following calculations are made:

Feed rate of Methane Feed(kg/h) 42830.00

Density of Liquid Methane (kg/m3) 422.62

Volume of Liquid Methane (m3/h) 101.34

Total volume stored (m3) 7296.77

Table 9-23: Calculations for storage volume

Total volume of liquid methane (natural gas) to be stored = 7296.77m3 (3 days’ worth)

20% volume allowance has been made for the vaporization of liquid methane within tanks, as

well as the innage/outage of the tanks.

Total volume needed = 7296.77 x 1.2

= 8756.12m3

Designing each tank to have a dimension of 55m (Diameter) by 35m (Height),

volume of each tank = 3022.25m3

Hence 3 tanks are necessary.



A proposed plot plan of the storage tanks is as follows:

Fig 9-4: Plot plan of storage tanks for natural gas feed

9.6.1 Capital Costs

Initial capital cost will constitute the 3 tanks as well as piping, heater and cooler.

Based on CAPCOST,

Cost for 1 tank = US$431,000

Cost for 3 tanks = US$431,000 x 3

= US$1,293,000

Cost for piping, heater and cooler has been taken to be 20% of tank costs.

Hence their cost = US$1,293,000 x 20%

= US$258,600

Total capital cost = US$1,293,000 + 258,600

= US$1,551,600



9.6.2. Operating Costs

Land Area needed (m3) Unit cost for land (US$/m3.yr) Land Cost (US$/yr)

21,025 8.84 185,829

Table 9-24: Operating Costs

Total duty (MWh) Unit cost for electricity (US$/MWh) Total Electricity cost (US$/yr)

885 100 88,520

Table 9-25: Electricity Costs

9.6.3 Overall Costs

Cost for the 1st year constitutes the initial capital cost in addition to land rental cost and

utility (electricity) cost. The remaining years in the course of plant operation will only involve

the operation costs, i.e. land rental costs and utility cost.

Assuming a storage facility is erected together with the construction of the plant, starting

from Year 0 that is, the cost breakdown is as follows:

Cost for 1st year 1,825,949 US$

Cost for 2nd year 274,349 US$

Cost for 3rd year 274,349 US$

Cost for 4th year 274,349 US$












Total Cost for all 15 years 5,666,836 US$

Table 9-26: Costs for storage facilities over 15 years



Making a conservative allowance of 40% to take care of any unforeseen expenses,

total cost of such a storage facility = US$5,666,836 x 1.4

= US$7,933,570

From Section 9.3.3, our fixed capital investment(FCI) for the main plant has been

determined to be at US$147,474,319.

If we were to build a storage facility to last for 15 years, it will constitute only

(7,933, 570/147,474,319) x 100%

= 5.4% of FCI

From an economic point of view, 5.4% is a fairly small percentage. From this cost

analysis, it can be seen that erection of the storage facility can be done at a cost which is only

around 1/20 times of the capital investment made into the parent steam-methane reforming plant,

rendering the former feasible as a project. As mentioned, the presence of such a storage facility

for stand-by feed will to some extent, provide the much-desired security if the plant management

is concerned with a steady provision of natural gas.

Similar to the main plant, a SHE analysis on the installation of storage tanks will have to

be conducted. This is especially necessary as the tank contents involve methane which is a highly

flammable substance. In view of the potential hazards associated with the feed, it is

recommended that the storage facility be sited at a safe distance away from the main steam-

reforming plant.

9.6.4 Economic Compensation

Despite the attractiveness of a feed-storage, for our design we have assumed provision of

natural gas via piping with no failures on any day of the year.



To back the decision of employing no storage tanks in our design, we have made the

following important assumption:

Should there be any failure in feed-delivery, compensations can be sought from the vendor

This is in view of the potential losses that can result from the inavailability of the natural

gas feed. Such compensations due to disruptions in service should have already been pre-agreed

upon as part of the terms of contracts between the plant management and their vendors.

However, the economic losses that arise from such a undesirable scenario can be difficult

to quantify sometimes. Plant management may consider employing the services of the financial

specialists to come up with a closer estimation to the actual damages suffered from a disruption

in feed supplies.

9.7 RECOMMENDATIONS

Based on the preliminary economic analysis, in order for the plant to break even, the

plant has to sell its hydrogen at US$2.43/kg. This analysis assumes a required rate of return of

10%, 18% corporate tax on income, 15 years discounted payback period and that all products of

this steam-reforming operation will be able to be sold in the market.

A more detailed analysis will encompass the inflation factor, as it is expected that the

company will have to increase the price of hydrogen over the years, following the expected

increase in operating costs, i.e. cost of utilities as a result of increases in crude oil prices,

increased rent cost etc. Pricing of the hydrogen heavily depends on the pricing strategies of other

competitors in the market, and in the event the minimum price which hydrogen should be priced

at (i.e. US$2.43/kg for our case) is significantly lower than the market average, say US$2.80/kg,

the company may want to sell its hydrogen at a higher price. Every cent increase in hydrogen

prices generates corresponding increase in profits as well as reduces the number of years which



the company needs to break even with the initial capital cost. Taking for example, if the

company chooses to price its hydrogen at US$2.70/kg, the company may actually require only 7

years for cumulative discounted cash flow to reach zero, as opposed to the 15 years necessary if

hydrogen is priced at US$2.43/kg. Hence, we can see that pricing of hydrogen depends a great

deal on the strategy which the company would like to employ:

1. Does the company want to gain substantial market share in the market by selling at a

rate way below market average?

2. Does the company want to recover all its investments in a shorter period by selling

hydrogen at a higher price, while still maintaining below market price?

A way to reduce the cost of manufacture, so as to increase the profit margins, is to reduce

the utility costs. Cogeneration facilities can be implemented so as to be self sufficient in the

production of electricity. However, the feasibility of this recommendation has to be determined

through detailed calculations of the fixed capital investment costs of these facilities.

On computations of the equipment costs, different vendors may have different quotes for

the required equipments. Therefore, rather than relying on simplified correlations (such as using

CAPCOST) to find the costs, it may be better to refer to the available catalogue from the vendors

for more current and realistic cost estimations. This is made possible with the advanced

telecommunication and readily available information from vendors working with firm’s parent

company.

Assuming that the designed plant is already in operation, in the event that all the above

recommendations do not lead to an accurate economic analysis that result in a positive NPV, the

firm can consider forming strategic alliance with some other firms that have prior experiences in

building similar plants, to reap the internal and external economies of scales possible with such

an alliance. Alternatively, the company may consider approaching Chemical Engineering

Consultancies for suggested improvements on the plant designs and operation procedures.



Financial or risk-management consultancies may be approached for a better estimation of present

and future financial status for the company. Should there be a definite forecast of unhealthy

financial status for the company in time to come, a re-assessment of the pricing strategies and

operation contingencies may be crucial.

9.8 CONCLUSION

Based on the assumptions made, the preliminary economic analysis shows that it is

possible to build a profitable steam-methane reforming plant. If we are concerned with a required

rate of return of 10% where payback period is 15 years, we can price our hydrogen at

US$2.43/kg, which is close to the suggested range ($1.90/kg to $2.30/kg) of hydrogen prices

provided by NETL (National Energy Technology Laboratory) [3]. Our price is obtained using

the Goal-seek function under Microsoft Excel by equating Cumulated Discounted Cash Flow in

the 15th year to zero. Any pricing of hydrogen above this rate increases our profit margins,

reducing the number the years which full returns can be realized. Again, here we are making the

assumption that any hydrogen produced will be able to find its customer in the market.

Competitiveness of the price, as mentioned, depends largely on the pricing strategies of other

hydrogen vendors in the market.

Factors that will make this economic analysis a more realistic study have been mentioned

under the recommendation section. As highlighted earlier, making an accurate assessment of

these factors in our work is beyond our expertise and hence employing finance specialists to

perform these complicated tasks will be the recommended option. In addition, chemical

engineering consultants can be approached for suggestions and further optimization on the

overall design and operations. In any event of a disruption in natural gas provision via piping,

compensations can be sought from the vendors, and availability of a legal support may reduce

the complications that arise from such a scenario.



REFERENCES

[R1]: Turton R., Bailie R.C., Whiting W.B. and Shaeiwitz J.A., “Analysis, Synthesis, and Design

of Chemical Process”, Prentice Hall, 2003

[R2]: Inland Revenue Authority of Singapore, http://www.iras.gov.sg/, retrieved on 12 April

2008.

[R3]: National Energy Technology Laboratory ,

http://www.netl.doe.gov/technologies/hydrogen_clean_fuels/systems_studies.html, retrieved on

12 April 2008

[R4]: Guthrie, K. M., “Data and Techniques for Preliminary Capital Cost Estimating”, Chem.

Eng., 1969

[R5]: Perry, R.H., and Green, D.W., “Perry’s Chemical Engineers’ Handbook”, 7th Edition,

McGraw Hill, New York, 1997.

[R6]: Peters, M. S., Timmerhaus, K. D., “Plant Design and Economics for Chemical

Engineers” McGraw Hill, USA, 1968.

[R7]: J.R. Couper et al., “Chemical Process Equipment: Selection and Design”, Elsevier, 2005.

[R8]: R.K. Sinot, “Coulson and Richardson’s Chemical Engineering”, Vol. 6, 4th Edition,

Oxford 2005.

[R9]: Jurong Town Council. JTC's Land Rents and Prices, Retrieved on 2 April 2008 at

http://www.jtc.gov.sg/products/land/industrialland/pages/index.aspx

CN 4120: Design II Team 32 Safety, Health & Environment (S.H.E.)


Chapter 10 : SAFETY, HEALTH & ENVIRONMENT (S.H.E.)

10.1 INTRODUCTION

Safety is of paramount importance in the operation of any chemical facility. Many

industrial accidents, such as the infamous Bhopal incident, serve as poignant reminders of the

significance of safety measures in preventing the escalation of minor accidents into major

catastrophes. A safe plant would make both economical sense, and uphold a positive image of

the company as a responsible corporate citizen in this global new economy.

To cover all safety aspects of plant operation, the following studies have been carried out:

1. Hazard and Operability review on the operations of the Furnace/SMR integrated system

2. Plant layout of the plant to ascertain preliminary site area and required safety clearances

3. Possible issues related to the occupational safety and health of plant personnel

Furthermore, the possible impact that the operation of a hydrogen plant has on the

environment would also be investigated in this report. It is essential that possible pollutants are

minimized and mitigation measures are duly mapped out in the preliminary design of a chemical

facility. This would be effected with the implementation of a risk assessment matrix.

A product life cycle assessment on hydrogen would also be used as a tool to further

illustrate the environmental impact of the manufacture of hydrogen through the steam methane

reforming route.



10.2 HAZARDS AND OPERABILITY STUDIES (HAZOP) REVIEW

A HAZOP study is used to identify the potential hazards in a chemical facility and is

generally an effective method to determine operational hazards in any facility [R1]. In this

preliminary design of a hydrogen plant, it was identified that the safe and functional operation of

the furnace and the SMR unit were crucial to the smooth running of subsequent downstream

units.

A key issue that involves the integrated furnace/SMR system would be the possibility of

explosion. This could arise because the integrated furnace/SMR system operates at the highest

temperature and pressure in the plant. Furthermore, the presence of volatile fuel gas/air mixtures,

coupled with methane from the natural gas feed and hydrogen formed after reaction in the SMR

unit could also be possible explosive hazards.

Hence a HAZOP study was employed to further investigate and analyze any potential

impediments to their safe operation. This was done with reference to the P&ID diagram

presented in the Chapter 11 on Instrumentation and Control.



Project Name: Production of Hydrogen via Syngas Route Process: Steam Methane Reforming Section: Steam Methane Reformer Reference Diagram: Furnace & SMR P&ID Table 10-1: HAZOP study for SMR preheated reactor feed

Study Node

Process Parameters

Item Deviations Possible Causes Possible Consequences Safeguards Actions Required

Preheated Reactor Feed

Flow 1A No 1. Control valves, FCV-101, FCV-102 fail close

2. Controllers, FC-101, FC-102 fail and close valves

3. Three-way valves fails close

4. Plugging of upstream piping

5. Plugging in expander E-101 and E-102

6. Complete plugging of heat exchangers HX-101 and HX-102

7. Operators’ error in isolating pipes for maintenance

1. No SMR reaction 2. Shutdown of

downstream units 3. Increase in

temperature of flue gas

4. Disrupt preheating of other process streams in convection section

1. Installation of low flow alarm

2. Installation of redundant expander E-102

1. Regular operators’ training

2. Review of operating procedures for proper pipeline isolation

3. Regular inspections and maintenance of expander E-101 and E-102, heat exchangers HX-101 and HX-102, piping, valves and fittings

1B

High 1. Control valve, FCV-101, FCV-102 fail open

2. Controller, FC-101, FC-102 fail and open valves

3. Flow indicator fails, indicating low

4. Error in flow control ratio

1. Increase in reactor pressure

2. Rupture of SMR tubes, leading to furnace explosion

3. Lower conversion due to lowered residence time

4. Increased corrosion of inner SMR tube

1. Installation of safety valves PSV-001, 002 and 005

2. Flow indicators FI-203 for flow monitoring


2. Regular inspections and maintenance of pipings, valves and fittings



5. Operator error in valve control

6. Faulty safety valve, fails to open

surface due to increased flow velocity


1C Low 1. Control valves FCV-101, FCV-102 fail partially open

2. Controllers FC-101, FC-102 fail and open valve partially

3. Flow indicator FT-101 fails, indicating high

4. Plugging of heat exchangers HX-101 and 102

5. Operator error in valve control

6. Safety valves PSV-001,002 and 005 not tightly close

1. Higher conversion rate due to longer residence time

2. Lower throughput 3. Possible

sedimentation of contaminants in pipeline

1. Installation of low flow alarm FLA-203

1. See 1B

1D Reverse 1. Pressure build-up in SMR

1. Hydrogen embrittlement in feed pipeline not designated to carry hydrogen

2. Damage to expanders E-101 and 102

1. Installation of check valves

2. Flow indicator FI-203 for flow monitoring

3. Installation of safety valve PSV-005 at the entrance to SMR

1. Decrease the furnace firing rate

Pressure 1E High 1. Coking of catalyst 2. See 1B

1. Increase in reactor pressure


1. Installation of safety valve

2. Installation of PI-101 and PI-102 to check for catalyst abnormalities through higher than normal pressure

1. Regular inspections and maintenance of valves, pipings, fittings and catalyst

2. Review of current operating



drop procedures to check for contributory factors to coking

3. Consider using medium pressure steam to decoke catalyst for maintenance

1F Low 1. See 1C

1. See 1C

1. See 1C

1. See 1B

Temperature 1G High 1. Control valve fails, leading to increased flow in the heating stream in HX-102

2. Inefficient expander E-101 and 102, leading to under expansion

3. Increased preheating of natural gas feed

1. Higher rate of catalyst coking

2. Shorter catalyst life, leading premature changeout

1. Installation of temperature indicator TT-101

2. Installation of flow control valve FCV-103 on the heating stream

1. Regular inspections and maintenance of valve and expander

2. Increase in generation of HP steam, which effectively reduces preheating of natural gas feed

1H Low 1. Control valve FCV-103fails, leading to decreased flow in the heating stream in HX-102

2. Poor piping insulation 3. Shell-side fouling of

heat exchanger 4. Inadequate preheating

of natural gas feed

1. Lower conversion in SMR tubes with the maintenance of furnace duty

2. Higher furnace duty required, which reduces the operating lifespan of the furnace

3. Increase usage of fuel gas feed

1. Installation of temperature indicator TT-101

2. Installation of flow control valve FCV-103 on the heating stream

1. Regular inspections and maintenance of valve and heat exchangers, HX-102

2. Decrease in generation of HP steam, which effectively increases preheating of natural gas feed



Composition

of natural

gas

1I High 1. Wrong setting of flow control ratio due to operators’ error

2. Controllers, FC-101, FC-102 fail

3. Control valves FCV-101, FCV-102 fail

1. Increase in catalyst coking due to reduced steam

1. Installation of controllers FC-101, FC-102

2. Installation of control valves FCV-101, FCV-102

1. Regular inspection and maintenance of valves and controllers


1J Low 1. See 1I

1. Higher cost due to large amounts of steam being injected

1. See 1I

1. See 1I



Table 10-2: HAZOP study for SMR reactor effluent Study Node

Process Parameters


Reactor Effluent

Flow 2A No 1. Plugging of reactor outlet

2. See 1A 3. Plugging of catalyst

1. Shutdown of downstream units



2. Regular inspections and maintenance of catalyst, pipings, valves and fittings

2B

Low 1. Partial plugging of reactor outlet

2. See 1B

1. See 1C



Pressure 2C Low 1. High pressure drop along tubes

2. Coking of catalyst

1. See 1H

1. Installation of PI-101 and PI-102 to check for catalyst abnormalities through higher than normal pressure drop

1. Review of current operating procedures to check for contributory factors to coking

2. Consider using medium pressure steam to decoke catalyst for maintenance

Temperature 2D High 1. Increased furnace firing

1. Coking of pipelines and downstream units

2. Corrosion of pipes due to high temperature

1. Installation of temperature indicator, TI-106

1. Reduce furnace firing

2E Low 1. Decreased furnace firing

1. Low conversion of natural gas to hydrogen

1. See 2D 1. Increase furnace firing

Composition 2F High 1. Low reaction 1. Eventual product off 1. Installation of 1. Monitoring of



of natural

gas temperature

2. High flow rate, leading to shorter residence time

3. Possible deactivation of catalyst

specification

methane analyzer A-004

2. Installation of temperature indicator TI-106

temperature and adopting necessary rectifications

2. Consider using medium pressure steam to decoke catalyst as regular maintenance

3. Catalyst changeout to be carried out if necessary



Project Name: Production of Hydrogen via Syngas Route Process: Combustion of Fuel Gas Section: Furnace (Radiant Section) Reference Diagram: Furnace & SMR P&ID

Table 10-3: HAZOP study for furnace (radiant section) fuel gas feed Study Node

Process Parameters


Fuel Gas Feed

Flow 3A No 1. Control valve fails close

2. Controller fails and closes valve

3. Complete plugging of fuel gas feed line

4. Operators’ error in isolating pipe for maintenance

5. Rupture of feed line 6. Opened safety valve

PSV-012. All flow redirect to safety valve

1. Flame extinguished due to lack of fuel supply

2. Pressure build-up upstream of blockage / closure

3. Little or no conversion attained in SMR

4. FCV-108 will close and lead to pressure buildup on preheat air line

1. Availability of make-up fuel gas to ensure continuous supply


1. Regular inspections and maintenance on valves, pipings and fittings

2. Regular operator training

3. Review standard operating procedure for pipeline isolation

4. Shut down of air blower B-101 in the event of pressure build-up

5. Adopt evacuation procedures for fuel gas release

3B

High 1. Control valve (Main Fuel gas line) fails open

2. Controller FIC-05 fails and opens valve

3. Error in flow meter FT-105 readings; indicating low

4. Control valve fails

1. Increase in temperature of flue gas

2. Thermal creeping of furnace tubes and walls

3. Unneeded usage of additional fuel gas and air

1. PSV-012 to relieve any excess gas to flare

2. Control valve

1. Regular inspections and maintenance of valves, indicators, pipings and fittings

2. Shutdown of fuel gas to prevent pipe rupture



(Make-up CH4) open 5. Faulty safety valve

PSV-012, fails to open

4. Rupture of pipeline due to high pressure

5. Flame impingement of tubes due to unstable flame

3. Adopt emergency shutdown procedures


3C Low 1. Control valve fails partially open / close

2. Controller FIC- 05 fails and opens / closes partially

3. Partial blockage / leakage of fuel gas feed line

4. Error in flow meter FT-105 readings; indicating high

5. Safety valve PSV-012 not tightly closed

1. Insufficient flame height due to low fuel supply

2. Optimal SMR conversion not achieved


4. Sedimentation of contaminants along the piping due to low fuel gas flow

1. Availability of make-up fuel gas to ensure continuous supply


1. Regular inspections and maintenance on valves, pipings, indicators and fittings


3D Reverse 1. Pressure build-up in furnace

1. Occurrences of backfiring

1. Installation of check valve

2. Installation of flame arrestor

1. Regular inspections and maintenance of check valve and flame arrestor

Pressure 3E High 1. Control valve fails open

2. Controller FIC-05 fails and opens valve

3. Pressure indicator fails, indicating low

4. Faulty safety valve PSV-012, fails to open

1. Increase in temperature of flue gas

2. Thermal creeping of furnace tubes and walls

3. Unneeded usage of additional fuel gas and air

4. Rupture of feed line due to overpressure

5. Excess flow to

1. PSV-012 to relief any excess gas to flare

1. See 3B



furnace resulting in unstable flame and flame impingement

3F Low 1. Control valve fails partially open / close

2. Controller FIC-05 fails and opens / closes partially

3. Partial blockage / leakage of fuel gas feed line

4. Error in flow meter FT-105 readings; indicating high

5. Safety valve PSV-012not tightly closed

1. See 3C 2. Occurrences of

backfiring

1. See 3C

1. See 3C

Temperature 3G Low 1. PSA malfunction 1. Lower furnace efficiency

1. Install temperature indicator

Composition

of H2

3H High 1. Deactivation of catalyst in PSA

1. Unstable flame leading to possible flame impingement

1. Installation of safety valve PSV-012

1. Changeout of catalyst in PSA if necessary

3I Low 1. PSA purge too low 1. Insufficient flame height due to low fuel supply


1. Introduction of makeup fuel gas



Table 10-4: HAZOP study for furnace (radiant section) preheated air feed Study Node

Process Parameters


Preheated Air Feed

Flow 4A No 1. Control valves fails close

2. Controller FIC-107 fails and closes valves

3. Loss of air supply due to damaged air blowers, B-101, B-102

4. Complete blockage of air feed line

5. Operators’ error in isolating pipe for maintenance

6. Rupture of air feed line

1. Flame extinguished due to lack of oxygen supply


3. Unneeded usage of fuel gas

4. Little or no SMR conversion attained

5. Fuel gas released to the environment due to incomplete combustion

1. FI-108 to indicate flow

2. FCV-108 to control flow



2. Regular inspections and maintenance of blower



4B High 1. Control valve fails open

2. Controller FIC-107 fails and opens control valves


4. Operators’ error 5. Error in flow meter

FT-107 and 202 readings; indicating low

1. Excessive flame height

2. Pressure build-up in furnace





4C Low 1. Partial plugging of control valves

2. Partial loss of air

1. Fuel gas released to the environment due to incomplete


2. FCV-108 to control

1. Regular inspections and maintenance of



supply due to damaged blower

3. Control valve FCV-107 fails to respond

4. Controller FIC-107 fails to respond

5. Piping or fitting leakage / partially blocked


7. Operators’ error 8. Error in flow meter

FT-107 and 202 readings; indicating high

combustion 2. Optimal SMR

conversion not achieved


flow 3. Installation of low

flow alarm FLA-201

pipings, valves and fittings


4D Reverse 1. Pressure build-up in furnace

1. Occurrences of backfire

1. Installation of check valve


1. Regular inspections and maintenance of check valve and flame arrestor

Pressure 4E High 1. CV fails open 2. Controller FIC-107

fails and opens control valves

3. Pressure indicator fails

4. Blower E-101 and 102 operating higher than normal

1. Pressure buildup within piping, leading to rupture

2. Excess flow to furnace leading to unstable flame

1. Installation of relief valve



1. Regular inspections and maintenance of valves, indicators, and blower

4F Low 1. See 4C 2. See 4E.2


2. Insufficient flame height or extinguished flame

1. Install low flow alarm FLA-202






Table 10-5: HAZOP study for furnace (radiant section) flue gas effluent Study Node

Process Parameters

Item Deviations Possible Causes Possible Consequences Safeguard Actions Required

Flue Gas Effluent

Flow 5A No 1. Malfunction of all burners

2. Complete blockages of fuel gas and air feeds

1. Reduced steam generation

2. Reduced preheating of natural gas

3. 3. Reduced preheating of air

1. Installation of peek holes

1. Regular checks via peek holes on burning

2. Regular inspection and maintenance of piping for fuel gas & air feeds

5B

High 1. Higher than normal fuel gas feed

2. Higher than normal air feed

1. Rupture of convection section due to pressure build-up

2. Collapse of tubes in convection section

1. Control valves that control fuel gas and air feeds

2. Lower feed rates fuel gas and air

5C

Low 1. Malfunction of burner(s)

2. Malfunction of ignition source

1. Reduced steam generation

2. Reduced preheating of natural gas

3. 3. Reduced preheating of air

1. Installation of peek holes


2. Regular inspection and maintenance of piping for fuel gas and air feeds

Pressure 5D High 1. See 5B 1. See 5B 1. See 5B 1. See 5B

5E Low 1. See 5C 1. See 5C 1. See 5C 1. See 5C

Temperature 5F High 1. High air / fuel gas flow rates

2. Inefficient heat transfer due to SMR tube degradation and fouling

3. High temperature of combustible air feed

1. Overheating and weakening of tubes in convection section

2. Cracking of furnace walls

3. Formation and emission of NOx due to higher temperature within furnace

4. Increased preheating

1. Installation of control valves that control fuel gas and air feeds

1. Lower feed rates of fuel gas and air

2. Lower combustible air temperature



of other process streams in convection section, leading to process disturbances in other units

5G Low 1. Low air / fuel gas flow rate

2. Inefficient furnace burning

3. Low temperature of combustible air

1. Less steam generation and preheating of other process streams in convection section, leading to process disturbances in other units

1. Installation of control valves that control fuel gas and air feeds


2. Regular inspection and maintenance of piping for fuel gas and air feeds

3. Increase temperature of combustible air

Composition

of CO

5H High 1. Incomplete combustion of fuel gas

1. Furnace flooding leading to afterburn in convection section and possible explosion.

2. Emission of CO and fuel gas to the atmosphere

1. Installation of CO and O2 analyzer in the convection section.

1. Lower fuel flow rate.

2. Slowly increase combustible air flow rate.



Project Name: Production of Hydrogen via Syngas Route Process: Heat Exchange of Flue Gas Section: Furnace (Convection Section) Reference Diagram: Furnace and SMR P&ID

Table 10-6: HAZOP study for furnace (convection section) condensate

Study Node

Process Parameters


Condensate

Flow 6A No 1. Control valve FCV-104 fails close

2. Controller FC-104 fails

3. Bypass valve fails close

4. Low temperature reading on temperature indicator TE-104

5. Operator error in isolation of piping for maintenance

1. No steam generation to buffer the flue gas temperature fluctuations

2. Too much preheating of other process streams in convection section, leading to disturbances in the other process units

3. Possible creeping of steam generation tube in convection section

4. Stack gas temperature would be elevated

1. Installation of bypass valve

2. Installation of flow controller FC-104 and valve FCV-104

1. Regular inspection and maintenance of valves, controllers and indicators


3. Review of current operating procedures

6B High 1. Control valve FCV-104 fails open


3. Bypass valve fails open

4. High temperature reading on temperature indicator TE-104

1. Rupture of steam generation tube in convection section

2. Too little preheating of other process streams in convection section, leading to disturbances in the other process units

3. Stack gas temperature would be

1. See 6A 1. See 6A



decreased, leading to acid corrosion issue in the stack

6C Low 1. Control valve FCV-104 fails partially close


3. Bypass valve fails partially close



1. Little steam generation to buffer the flue gas temperature fluctuations




1. See 6A 1. See 6A

Pressure 6E High 1. Control valve FCV-104 fails open


3. Bypass valve fails open

4. High temperature reading on temperature indicator TE-104

1. Rupture of steam generation tube in convection section

2. Too little preheating of other process streams in convection section, leading to disturbances in the other process units

3. Stack gas temperature would be decreased, leading to acid corrosion issue in the stack

1. See 6A 1. See 6A

6F Low 1. Control valve FCV-104 fails partially close

1. Little steam generation to buffer the flue gas

1. See 6A 1. See 6A




3. Bypass valve fails partially close



temperature fluctuations






Table 10-7: HAZOP study for furnace (convection section) natural gas feed

Study Node

Process Parameters


Natural Gas Feed

Flow 7A No 1. Disruption of supply 2. Pipeline rupture

1. Discharge of natural gas into environment

2. No reaction in SMR 3. Higher preheating of

combustible air 4. Elevation of stack

gas temperature

1. Ensure duplicity in supplies

1. Regular inspection and maintenance of supply pipelines

2. Increase the amount of steam generation to offset the increase in preheating of combustible air and stack gas temperature

7B

High 1. Surges in supply 1. Rupture of natural gas tubes in convection section, leading to possible furnace explosion

2. Less preheating of combustible air


1. Consider the possible installation of feed surge drum to regulate the flow of natural gas and eliminate any fluctuations

7C Low 1. Fluctuations in natural gas feed flow

2. See 7A

1. See 7A.1 2. Lower SMR

throughput

1. See 7A 1. See 7A

Pressure 7E High 1. See 7B

1. See 7B

1. See 7B

1. See 7B

7F Low 1. See 7C

1. See 7C

1. See 7C

1. See 7C



Table 10-8: HAZOP study for furnace (convection section) combustible air feed

Study Node

Process Parameters


Combustible Air

Flow 8A No 1. Loss of air supply due to damaged air blowers, B-101

2. Complete blockage of air feed line

3. Complete plugging of air filter

1. Elevation of stack gas temperature

2. Flame extinguished due to lack of oxygen supply

3. Little or no SMR conversion attained

1. Availability of redundant blower, B-102

1. Periodic cleaning or replacement of air filter

2. Regular inspections and maintenance of air blowers

8B

High 1. Accidental operation of two blowers B-101 and 102

2. Safety valve PSV-010 fails to open

1. Decrease in stack temperature, leading to possible acid corrosion in the stack

2. Rupture of air pipelines in the convection section, leading to possible ignition of residual fuel gas



2. Review of blower operating procedures

8C Low 1. Damaged blower, B-101 resulting in reduced air flow

2. Partial blockage of air filter

3. Partial plugging of air feed line


2. Elevation of stack gas temperature

1. Availability of redundant blower, B-102

1. See 8A

Pressure 8E High 1. See 8B

1. See 8B

1. See 8B

1. See 8B

8F Low 1. See 8C

1. See 8C

1. See 8C

1. See 8C



Table 10-9: HAZOP study for furnace (convection section) stack gas

Study Node

Process Parameters


Stack Gas

Temperature 1A Low 1. Too much heat transfer to process and utility streams in convection section

2. Low flue gas temperature from radiant section

1. Condensation of acid gases, leading to corrosion in stack

1. Installation of temperature indicator, TI-105

1. Lower rate of steam generation

2. Increase flue gas temperature by increasing rate of burning



Project Name: Production of Hydrogen via Syngas Route Process: Heat Exchange of Flue Gas Section: Furnace Reference Diagram: Furnace & SMR P&ID

Table 10-10: HAZOP study for furnace

Study Node

Process Parameters


Furnace Pressure 10A High 1. Faulty control system, causing stack damper to remain closed

2. Increase in rate of burning


2. Cracking of furnace walls



2. Installation of high pressure alarm

1. Regular inspections and maintenance of control systems, tubes and damper

2. Review of current operating procedures if high pressure occurrences are frequent

Temperature 10B High 1. Too much fuel gas and air feed

2. No reactant flow in SMR tubes

1. Thermal creep of SMR tubes

2. Furnace walls would crack

1. Installation of temperature indicator TI-202, TI-203

2. Install control valve FCV-108

1. Regular inspections and maintenance of valve and control system



10.3 PLANT LAYOUT

Plant layout is the spatial arrangement of items of process vessels and equipment and

their connection by pipes, ducts, conveyors or vehicular transportation, as defined by

Mecklenburgh [R2]. It is essential that plant layout is carefully thought out to satisfy the following

key considerations:

Cost-effective use of space

Risk-free and efficient construction

Reliable, efficient and safe operations

Ease of maintenance and repair of process and associated auxiliary units, carried out on

site or ex situ

Minimal hazard and nuisance caused to the public

Besides safety and economic concerns, other factors such as process requirements, fire-fighting

and emergency capabilities, administrative and medical support infrastructure would also have to

be incorporated into the overall plant layout. An efficient transportation network within the plant

for the movement of materials, personnel and emergency services would also be needed. These

key issues would be duly covered in the following paragraphs.

10.3.1 Segregation

A main process area was designated, housing the primary process units such as the

furnace, SMR, HTS, LTS and the PSA unit. This facilitated the movement of material from one

process unit to the other and allowed for lower piping cost.

For safety and loss prevention, this main process area was deliberately sited away from

the main administrative building, laboratory and other buildings which house work personnel.

Furthermore, to control access to the main process area and also to the auxiliary units and waste

treatment area, a security fence was erected and an ‘official access only’ system was

implemented, in view of increasing security and intrusion concerns.



A safety clearance of at least 60 m would also be needed between this main process area

and the main buildings where plant personnel are housed [R2].

10.3.2 Transportation Considerations

A rectangular grid was employed in the preliminary plant layout as it is often the most

cost-effective arrangement. This is because overhead piperacks, sewer systems, underground

piping systems, trenches, electric cables and instrument lines very often follow the road layout

[R2]. Therefore, curved roads should generally be avoided to facilitate the laying of these lines and

systems.

Roads should also be built in such a way so as to allow easy access for emergency and

fire-fighting services. This can be accomplished with the use of a peripheral road system which

allows for at least 2 approaches to all major fire risks, e.g. the main process area. Roads in the

plant should also have access to the public road system at a minimum of 2 points. Moreover,

they should be sited at a reasonable distance from any process unit or building in order to

effectively carry out firefighting or administer emergency care. The actual distance is usually set

at around 18 to 45 m [R2].

The main perimeter road was set to be 10 m wide. This is to facilitate the bulk of daily

vehicular traffic. Primary and secondary access roads were designated to be 6 m and 3.5 m wide

respectively. Adequate parking space was also allocated for work personnel, visitors and

loading/unloading vehicles. These were sited away from wind-blown dusts and the main process

area for safety and security purposes, and suitably sized to prevent congestion during shift

changeover [R2].

10.3.3 Administration

Plant personnel with more general site responsibilities, e.g. human resource and IT

departments, should be sited in an administrative building located in a non-hazardous area away

from the main process units. This building should be sited upwind of possible fumes and



emission from the process units, and near the main entrance to the plant to facilitate evacuation.

An estimated 50 administrative employees was used as the benchmark in this preliminary design,

hence according to Mecklenburgh [R2], an area of 500 m2 would suffice for preliminary layout

planning.

10.3.4 Laboratory

An onsite laboratory is critical in the setup of any plant as it provides a readily-available

source of analytical information on the quality of intermediates and products. This allows

engineers and operators to effect any changes to the process parameters in order to correct any

deviations. An estimated 20 laboratory staff would suffice for this preliminary design; hence an

area of 400 m2 would be used [R2].

10.3.5 Workshop

A workshop is needed for mechanical repairs and maintenance to be carried out onsite.

This will help to defray costs related to delays in repair and transport elsewhere. An estimate of

30 people would be used; hence an area of 600 m2 would be set aside for such a workshop [R2].

10.3.6 Control Room

The location of a control room is based on normal operating requirements and the need

for protection during emergency situations. Operators would be more inclined to be on the plant

should the control room be nearer to the plant. As a result, they would be more likely to observe

any malfunctions and plant deviations, hence preventing the occurrence of a serious fault.

Instrument cables would also be shortened if the control room is sited nearer the plant. It has also

been reported that at distances greater than 35 m and particularly over 100 m, operators will tend

not to go into the plant, especially in inclement weather [R2].

Nevertheless, the shorter the distance the control room is to the plant, the greater the need

for the room to be reinforced for personnel protection, and hence a higher cost is needed.



Therefore, a suitable distance of about 30 m between the control room and the main process area

was chosen in this design [R2]. An estimated area of 200 m2 was used for the control room.

10.3.7 Transformer Substation

The substation is usually sited at the edge of the plant, in areas of the lowest electrical

hazard rating. This will ensure a continuous electrical supply to the entire plant, critical to its

daily functioning.

10.3.8 Emergency Services

Emergency services such as ambulance and fire stations need to be given rapid access to

the entire site, without causing any hazard to the existing plant traffic. Preferably, they should be

housed outside the main fence but close to the main plant entrance. A 500 m2 area would be

recommended for this hydrogen plant.

An area adjacent to the main administrative building would also be set aside as an

emergency assembly area. This would help in the accounting of plant personnel should an

emergency incident arise.

10.3.9 Amenities (Medical Centre and Canteen)

The medical centre is essential in administering emergency and daily medical care to the

plant employees. Since it is frequently used by staff from all over the site, it should be located in

a central, non-hazardous area, preferably grouped together with other amenities such as the

canteen. This will help to cut down on time spent traveling to these amenities. Also, it should be

sited upwind from drifting fumes and noise from process units.

It was hence decided that an area of 30 m2 would be used for the medical centre, able to

cater to about 200 employees and contractors. An area of 200 m2 would be used in sizing the

canteen, also with the capacity to accommodate 200 personnel at one time [R2].



10.3.10 Process and Auxiliary Units

10.3.10.1 Furnace (Housing SMR)

The foremost considerations in the location of a furnace are safety issues since it is a

constant source of ignition. Where practical, furnaces should be sited in upwind locations so that

flammable gases or vapors are less likely to be blown towards the furnace, resulting in ignition.

Furnace transfer lines should also be kept short, with a common stack policy being employed.

Hence, process equipment which is directly connected to the furnace should be sited as close as

possible, without compromising on the recommended safety distance of 30 m from any

equipment which could be possible sources of ignition.2 Also, underground drain-points and

manhole covers within 30 m of the furnace walls should be sealed, and pits or trenches should

generally be avoided from extending under furnaces.

With regards to the abovementioned considerations, the furnace housing the SMR would

be located within the designated main process area housing the other major units. The

recommended safety spacing would be a distance of 15 m between the furnace and the HTS and

LTS reactors, and a distance of 30 m between the furnace and the PSA unit, which is considered

to be a process unit with a low flash point due to the high concentration of H2 present [R2].

10.3.10.2 Reactors (HTS, LTS), PSA and Knockout Drum

HTS and LTS are typical fixed-bed reactors loaded with catalyst in bulk between

supports within the reactor vessel. The PSA unit also faces similar considerations due to its usage

of bulk catalyst as an absorbent; hence it was mentioned together in this section.

As catalysts have a fixed effective shelf life of a few years, provisions have to be made

for the removal and loading of catalysts. The units have to be sufficiently elevated from the

ground so as to allow catalyst removal by mechanical transport, belt conveyors, fluid conveying

or hand trucks. Clearance should be allowed for the usage of mechanical drills and other

equipment should coking, sintering, or hard agglomeration of the catalyst [R2].



According to Mecklenburgh [R2], an area of 4 m2 must also be available at the base of

each reactor for the transport and temporary storage of fresh and spent catalyst.

It is also important to note that the PSA unit is used to concentrate the amount of H2.

With H2 being a vapor with a very low flash point of -273°C [R3], care would need to be exercised

in the safety distances needed. Since the PSA unit consists of 8 separate columns, it has been

recommended that such process equipment with low flash points should be spaced 2 m away

from one another [R2]. Also, it would be housed together in the main process area, at a distance of

at least 5 m from the HTS and LTS reactors, and a distance of at least 30 m from the furnace.

The knockout drum would be sited just next to the LTS unit.

10.3.10.3 Cooling Tower

Problems associated with the location of a cooling tower are often related to the large

volumes of very humid air which emanates from it, compounding the already high levels of

humidity normally experienced in Singapore. The moisture can lead to fog, precipitation and

corrosion issues in areas downwind of it [R2].

Towers should also be sited to mitigate the effects of wind drift on roads, rail, plant and

the neighborhood of the site. It is important to check that any possible corrosive emissions from

the vents of the HTS and LTS reactors, and the stack emissions from the furnace, would not be

entrained within the cooling tower.

Specifically for our project, a mechanical-draught tower was used. These require power

and may generate associated fan noise. If the entire plant was sited near a residential area, a

slower fan speed should be employed. Alternatively, buildings or sound screens could be erected

between the tower and the residential area.

Hence, the cooling tower has been separately sited far away from the main process area,

i.e. as an offsite facility, so as to prevent any entrainment of corrosive vapors.



It has been accorded a safety distance of at least 30 m from the process area, at least 60 m

from administrative buildings, at least 30 m from the fire station and at last 45 m from the main

plant substation [R4].

10.3.10.4 Heat Exchangers

Heat exchangers should be located within the conventional process unit plot area, in close

proximity to the equipment which they are associated with. This would minimize the cost of pipe

runs and also facilitate operator and maintenance access. Heat exchangers between two process

equipment which are far away should be sited at optimal points in relationship to pipe tracks [R2].

However, due to the close proximity of our major process units to one another, with the

exception of the cooling tower, it would not be a major issue for consideration in this preliminary

design.

If exchangers are located in pairs, or in larger groups, they can be stacked on top of one

another. A spacing of 0.45 m (18 inch) should be provided to allow maintenance to be carried

out on the flange bolts easily [R4]. By this means, there could be resultant savings on service

pipework, pipebridge and structural work etc [R2]. However, process piping and access steel work

may actually increase consequently; therefore a compromise has to be made. Generally, most

exchangers would also be placed on a base about 1 m above ground level for the provision of

drain connections.

As there are no restrictions on minimum safety clearances between heat exchangers and

the specific process units that are found in our preliminary design, it would be considered that

they occupy the space in between the safety clearances of the process units.

10.3.10.5 Flares

Flares are used to burn away excess gases in an emergency situation and also to flare

away off-specification gases such as H2. A sterile radius of at least 60 m should be maintained



around the flare, with only associated flare equipment and access roads allowed within this

radius. Also, the flare stacks should be located downwind from the main process areas and at

least 100 m away to provide for the dispersion of vapor releases [R2].

10.3.10.6 Wastewater Treatment Plant

A wastewater treatment plant is also needed to treat the condensate from the knockout

drum, the blowdown from the cooling water tower and any other liquid plant effluent before it is

discharged. It is usually sited at the perimeter of the site.

The estimated total effluent from the knockout drum and the cooling water tower is

around 873760 m3/year, assuming 8000 operating hours in a year. Comparing this with a water

treatment plant sited in Changi, Singapore5, which processes 292 million m3 of wastewater per

year (calculated from a daily rate of 800000 m3/day) and occupies an area of 55 ha (550000 m2),

the area needed for the onsite treatment plant would be estimated to be around 1650 m2.

Table 10-11: Base areas of process, auxiliary and other plant facilities

From Figure 10-1, the plant area was estimated to be around 29400 m2.

Process & Auxiliary Units Base Area

(m2)

Plant Facilities

Base Area (m2)

Furnace (housing SMR) 891 Administrative Building 500

HTS 9.4 Laboratory 400

LTS 8.6 Workshop 600

Knockout Drum 3.0 Control Room 200

PSA 144 Transformer Substation 160

Cooling Water Tower 733 Fire Station 500

Flare 25 Medical Centre 30

Wastewater Treatment Plant 1650 Canteen 200



Fig 10-1: Plant layout

10.4 OCCUPATIONAL SAFETY

10.4.1 Personal Protection Equipment (PPE)

The safety requirements for different tasks in different parts of the plant are different.

Thus it is important to consult the Safety, Health and Environment department to ensure that the

correct PPE are worn [R1] when performing each task.



Personal protection equipment is extremely important in preventing or reducing exposure

by providing a barrier between the worker and the workplace environment. This is for the benefit

of workers’ safety and health. However, by itself, it does not prevent accidents. It is only acting

as a secondary measure for safety and health purposes.

One of the most important equipment for this plant would be self-contained breathing

apparatus. This is because the dangers of this plant would mostly be caused by inhalation.

Therefore, this equipment has to be strategically placed in areas where leaks might occur,

especially in confined spaces. Other equipment needed would be safety glasses, safety gloves

and safety shoes.

Employers have to identify and provide the appropriate PPE for their employees. They

would also have to train employees in using and caring for the PPE. Periodically, they should

also review and check the PPE, replace worn out items and make the PPE program more

effective.

Employees should properly adhere to the regulations on wearing their PPE. They should

also attend training sessions for the use of their PPE. They should have the responsibility to take

care and maintain their PPE and inform their superiors if there is a need to replace these items.

10.4.2 Noise

Noise [R1] problems are commonplace in chemical plants. They are measured in decibels

(dB). According to regulations, noise levels should not exceed 85dB for an 8 hour workday.

Measures must be taken if the exposed noise level is higher than 85dB. Control of the noise can

happen at either the source or the receiver of the noise. At the source of noise, one can either

enclose the source using shields such as plywood or noise absorbing foams. It is also possible to

employ sound barriers to reduce the noise level transmitted to workers. Otherwise, at the receiver

level, ear plugs and ear muffs are the most commonly used PPE for regulating exposure to noise.



10.4.3 Ventilation

Ventilation [R1] is extremely important for this plant for the following reasons:

It can remove dangerous concentrations of flammable and toxic materials

It can be highly localized, reducing the quantity of air moved

Ventilation equipment is readily available and can be easily installed

It can be added to an existing facility

However, the major disadvantage is operating cost, which is rather prohibitive.

Ventilation systems comprise of fans and ducts that effectively dilute the contamination using

dilution ventilation. Fresh air is flowed in large amounts to dilute the contamination. Workers

would still be exposed to the contaminants, but in lesser concentrations.



10.5 OCCUPATIONAL HEALTH HAZARD IDENTIFICATION

It is important to identify any safety hazards brought about by the chemical nature of the reactants and products present in the

steam methane reformer. A good knowledge of the nature of these components is necessary for operators to know how to react to

exposure to these components.

Table 10-12: Summary of nature and associated hazards of chemicals [R6]

Chemical Threshold Limit Value (TLV)

Lower Exposure Limit and

Upper Exposure

Limit (LEL – UEL)

Personal Protective Equipment

(PPE)

Health Effects Acute Chronic

Methane N.A. 5 – 15% Safety glasses and/or face shields

Simple asphyxiant- reduces the amount of oxygen in the air. Exposure to oxygen-deficient atmospheres (less than 19.5 %) may produce dizziness, nausea, vomiting, loss of consciousness, and death. At very low oxygen concentrations (less than 12 %) unconsciousness and death may occur without warning. It should be noted that before suffocation could occur, the lower flammable limit for Methane in air will be exceeded; causing both an oxygen deficient and an explosive atmosphere.

N.A.

CO 25 ppm TWA

12.5 – 74% Safety glasses, safety gloves, safety shoes, and self-contained

Inhaled carbon monoxide binds with blood hemoglobin to form carboxyhemoglobin. Carboxyhemoglobin cannot take part in normal oxygen transport, greatly reducing the blood’s ability to transport oxygen. Depending on levels

N.A.



breathing apparatus during emergency

and duration of exposure, symptoms may include headache, dizziness, heart palpitations, weakness, confusion, nausea, and even convulsions, eventual unconsciousness and death.

H2 N.A 4% - 74.5% Wear chemical resistant gloves

Nausea, vomiting, difficulty breathing, irregular heartbeat, headache, fatigue, dizziness, disorientation, mood swings, tingling sensation, loss of coordination, suffocation, convulsions, unconsciousness, coma

N.A

CO2 5000 ppm TWA

Non-flammable

Safety glasses, safety gloves, safety shoes, and self-contained breathing apparatus during emergency

Carbon dioxide is a cerebral vasodilator. Inhaling large quantities causes rapid circulatory insufficiency leading to coma and death. Asphyxiation is likely to occur before the effects of carbon dioxide overexposure. Low concentrations cause increased respiration and headache. Product is a simple asphyxiant. Effects of oxygen deficiency may include any, all or none of the following: rapid breathing, diminished mental alertness, impaired muscle coordination, blurred speech, and fatigue. As asphyxiation progresses; nausea, vomiting, and loss of consciousness may occur, eventually leading to convulsions, coma and death.

N.A

These are the procedures to be taken for spills, fire-fighting and medical aid. Table 10-13: Summary of safety, fire-fighting and medical aid measures [R6]

Chemical Spill/leak measures Fire-fighting measures Medical Aid measures Methane Personal precautions: Wear self

contained breathing apparatus when entering area unless atmosphere is proved to be safe. Evacuate area. Ensure adequate air ventilation. Eliminate

Extremely flammable. Exposure to fire may cause containers to rupture/explode. If possible, stop flow of product. Move away from the container and cool with water from a

Inhalation: Remove victim to uncontaminated area wearing self-contained breathing apparatus. Keep victim warm and rested. Call a doctor. Apply artificial



ignition sources. Environmental precautions: Try to stop release. Clean up methods: Ventilate the area.

protected position. Do not extinguish a leaking gas flame unless absolutely necessary. Spontaneous explosive re-ignition may occur. Extinguish any other fire. In confined space, fire fighters must use self-contained breathing apparatus.

respiration if breathing stopped. Ingestion: Ingestion is not considered a potential route of exposure.

CO Personal precautions: Evacuate area. Eliminate ignition sources. Use self-contained breathing apparatus and chemically protective clothing. Ensure adequate air ventilation. Wear self contained breathing apparatus when entering area unless atmosphere is proved to be safe. Environmental precautions: Try to stop release. Clean up methods: Ventilate the area.

Extremely flammable. Exposure to fire may cause containers to rupture/explode. If possible, stop flow of product. Move away from the container and cool with water from a protected position. Do not extinguish a leaking gas flame unless absolutely necessary. Spontaneous explosive re-ignition may occur. Extinguish any other fire. Fire fighters must use self-contained breathing apparatus.

Inhalation: Remove victim to uncontaminated area wearing self-contained breathing apparatus. Keep victim warm and rested. Call a doctor. Apply artificial respiration if breathing stopped. Ingestion: Ingestion is not considered a potential route of exposure.

H2 Evacuate all personnel from affected area. Use appropriate protective equipment. If leak is in user’s equipment, be certain to purge piping with an inert gas before attempting repairs. If leak is in the container of container valve, contact closest supplier location.

If possible, stop the flow of hydrogen. Cool surrounding containers with water spray. Hydrogen burns with an almost invisible flame of relatively low thermal radiation. Hydrogen is very light and rises very rapidly in air. Should a hydrogen fire be extinguished and the flow of gas continue, increase ventilation to prevent an explosion hazard, particularly in the upper portions.

Inhalation: Remove victim to uncontaminated area wearing self-contained breathing apparatus. Keep victim warm and rested. Call a doctor. Apply artificial respiration if breathing stopped. Ingestion: Ingestion is not considered a potential route of exposure.

CO2 Personal precautions: Evacuate area. Non-flammable. If possible stop flow Inhalation: Remove victim to



Wear self contained breathing apparatus when entering the area unless atmosphere is proved to be safe. Ensure adequate air ventilation. Environmental precautions: Try to stop release. Prevent from entering sewers, basements and workpits, or any place where its accumulation can be dangerous. Clean up methods: Ventilate the area.

of product. Move away from the container and cool with water from a protected position. In confined space, fire-fighters must use self contained breathing apparatus.

uncontaminated area wearing self-contained breathing apparatus. Keep victim warm and rested. Call a doctor. Apply artificial respiration if breathing stopped. Skin/eye contact: Immediately flush eyes thoroughly with water for at least 15 minutes. In case of frostbite, spray with water for at least 15 minutes. Apply a sterile dressing. Obtain medical assistance. Ingestion: Ingestion is not considered a potential route of exposure.



10.6 ENVIRONMENTAL IMPACT ASSESSMENT

Environmental impact assessment (EIA) refers to the need to identify and predict the

impact on the environment and on man’s health and well-being of legislative proposals, policies,

programmes, projects and operational procedures, and to interpret and communicate information

about the impacts [R7]. Although there is no legal requirement for EIA to be carried out for

projects in Singapore [R8], an EIA would be carried out in this preliminary design of a hydrogen

plant so as to address the environmental aspects of undertaking the operation of a hydrogen plant.

10.6.1 Objectives

The primary objectives of an EIA are summarized as follows [R9]:

To determine methods to prevent or mitigate environmental damage

To mitigate environmental damage through the application of practical meditative actions

To make known to the public and the public or private bodies in charge of the project the

noteworthy environmental effects of such an undertaking

To make known to the public justifications of governmental approvals of undertaking

with substantial environmental effects

To encourage cross-agency interaction in the assessment of projects

To engage the public in the planning process

10.6.2 Risk Assessment Matrix

In order to assess the environmental impact of the daily activities of the hydrogen plant

and its accompanying wastewater treatment plant, a risk assessment matrix as shown in Table

10-14 was utilized to summarize the offending activities and its associated impacts on the

environment and work personnel. Criteria employed in this risk assessment matrix are reflected

in the accompanying legend, with risk ranking (RR) having considered the overall consequences

(OC) and the likelihood of occurrence (LH). The RR would serve as a clear indication of the risk

level of a particular activity in the operation of the hydrogen plant.



Table 10-14: Risk assessment matrix detailing the environmental impact of hydrogen plant operations

Activity/Process/ Product/ Services

Aspect/Associated Hazard

N/A/E Impact/Effect P PD ENV REP OC LH RR Mitigation Measures

Flaring 1. Complete combustion yields CO2, a major greenhouse gas

N 1. Increased greenhouse effect

IV IV III YES IV A M 1. Reduce frequency of excess waste gas emission, emergency flaring and off-spec flaring incidents

2. Try to achieve absolute combustion through use of excess air

3. Flare got opacity analyzer

4. Installation of air scrubbers

5. Periodic monitoring to prevent excessive emissions

2. Incomplete combustion yields methane (from natural gas), H2 (off-spec product), soot and CO


2. Emission of H2, a highly flammable gas and an explosive hazard

3. Soot can worsen to regional haze problem

4. CO can indirectly raise methane and tropospheric ozone, which also contribute to

III III II YES III B M



the global warming

Emission of stack gases during furnace operation

1. Emission of greenhouse gases such as CO2


IV IV III YES IV A M 1. Ensure efficient furnace operation, using the least possible fuel gas

2. Emission of NOx due to high temperature in furnace radiant section

N 1. Acid rain formation

III III II YES II A H 1. Install temperature sensor to monitor radiant section temperature

2. Install NOx controller

3. Emission of soot and CO due to incomplete combustion

N 1. Soot can contribute to regional haze problem

2. Emission of CO can indirectly

III III II YES III B M 1. Install opacity analyzer to detect excessive soot formation

2. Ensure excess air is present



raise methane and tropospheric ozone, which also contribute to the global warming

to minimize incomplete combustion

4. If sulfur is found in natural gas feed, SO2 could be released

A 1. Acid rain formation

III III II NO III D L 1. Use fuel gas feed with less sulfur

Vapor leakages from pipelines

1. Emission of methane, hydrogen, CO and CO2

A 1. Increased greenhouse effect due to methane, CO and CO2, leading to global warming

2. Emission of H2, a highly flammable gas and an explosive hazard

3. Exposure to CO, an asphyxiant, can result in dizziness,

II II III NO II C M 1. Isolation and repair of leaking pipes

2. Conduct regular checks



nausea, vomiting, loss of consciousness and death

2. Hot steam emissions due to ruptured pipe

A 1. Injury to work personnel

III IV IV YES IV C L 1. Conduct steam leak test

2. Clamp and repair leaking pipes

Storage of cooling water chemicals (biocides, corrosion inhibitor)

1. Leakage of non-volatile chemicals

A 1. Soil contamination

2. Water contamination

III IV I YES III C M 1. Set storage limits

2. Ensure proper storage of chemicals

3. Containment of leaks

4. Neutralization or dilution methods

5. Proper disposal of expired chemicals

6. Regular checks by cooling tower unit



operators 7. Periodic soil

quality tests

Knockout drum operation

3. Condensate leakage from drum and associated pipelines, carrying dissolved hydrocarbons and carboxylic acids)

A 1. Soil contamination

2. Water contamination

III IV I YES III C M 1. Containment of leaks

2. Regular pipeline checks

3. Periodic maintenance of drum and pipelines

Cooling tower operation

1. Water droplets or water mist aerosol generated as drift

A 1. Legionnaire’s Disease can be transmitted, leading to pneumonia in serious cases [R10]

II IV III YES III C M 1. Periodic disinfection of cooling tower using chlorine [R11]

2. Installation of drift eliminators

3. Usage of biocides to prevent accumulation of algae and scaling

2. Discharge of heated cooling water into surrounding water bodies

A 1. Raise temperatures of surrounding water

III IV I YES III D L 1. Ensure cooling water is not directly discharged into surrounding



bodies, decreasing amount of dissolved oxygen

water bodies 2. Installation of

temperature sensors at discharge outlet

Waste water treatment

1. Discharge of treated effluent

N 1. Water pollution

III IV IV YES IV A M 2. Installation of sensors for detection of elevated levels of chemical release

3. Monitor pH levels

Catalyst change out in reactors and PSA unit

1. Metal content present in catalyst

N 1. Soil contamination by metals

2. Water contamination by metals

III IV II NO III C M 1. Proper disposal of spent catalyst

2. Dust particles from catalyst fines

N 1. Affects respiratory system of work personnel

III IV III NO III C M 1. Don respiratory masks during catalyst change out operation



Legend for Table 10-14

Symbol Description

N / E / A Normal Operation / Emergency / Abnormal Operation

P Injury to People

PD Property Damage

ENV Environmental Impact

REP Repetitive

OC Overall Consequences

LH Likelihood of Occurrence

RR Risk Ranking

Probability Category Definition Probability A Possibility of repeated incidents A B C D E B Possibility of isolated incidents I C Possibility of occurring sometimes II D Not likely to occur III E Practically impossible IV

Consequence Category Considerations

Health/Safety Property Damage

Environmental Impact

I Fatalities / serious impact on public Large

community Major/Extended duration/Full scale response

II Serious injury to personnel / limited

impact on public Small

community Serious/Significant resource commitment

III Medical treatment for personnel /

No impact on public Minor Moderate/Limited response or short duration

IV Minor impact on personnel Minimal to

none Minor/Little or no response needed



10.6.3 Elements of Environmental Impact Assessment

10.6.3.1 Gaseous emissions

The air pollution caused by the H2 plant may include greenhouse gases such as carbon

dioxide, methane and NOx, hazardous gases such as carbon monoxide and sulphur dioxide, and

also flammable gases such as hydrogen. These gases could cause much harm to the environment,

for example, global warming due to greenhouse gases, acid rain due to acidic gases, ozone

depletion, and even fires and explosion caused by flammable gases. Some of these gases also can

cause health problems in people who are exposed to it without proper protection. Sulphur

dioxide, nitrogen dioxide and ozone in the lower atmosphere can cause respiratory diseases in

people. Emission limits of sulphur dioxide are 500mg/Nm3, and that of nitrogen dioxide are

700mg/Nm3.

Carbon monoxide is also dangerous and can cause death through asphyxiation in

excessive doses. Emission limits of carbon monoxide are 625mg/Nm3. During changeout of

catalyst, it could cause the broken, fine metal catalysts to emit into the air as respirable

suspended particles (PM10). PM10 refer to particulate matter of size 10mm and below. These

particles have health implications as they are able to penetrate into the deeper regions of the

respiratory tract. In very large amounts, the particles cause breathing and respiratory problems,

and aggravate existing respiratory and cardiovascular diseases. Hence, to reduce the emissions of

these vapours, mitigation measures as mentioned in Table 10-14 is essential to be adhered to.

10.6.3.2 Effluent discharge

The Pollution Control Department in Singapore regularly monitors water quality of

various inland water bodies and coastal areas. For this hydrogen plant, the effluent would come

from the cooling tower water and chemicals as well as the wastewater from the knockout drum

which comprises of hydrocarbons. The table below shows the allowable limits for trade effluent

discharged into a public sewer/watercourse/controlled watercourse.



Items of Analysis

Public Sewer Watercourse Controlled watercourse

Units in milligram per litre or otherwise stated

1. Temperature of discharge 45oC 45oC 45oC

2. BOD (5 days at 20oC) 400 50 20

3. COD 600 100 60

4. Hydrocarbon 60 10 -

5. Total Suspended solids 400 50 30

Table 10-15: Allowable limits for trade effluent discharge

Where, Biochemical Oxygen Demand (BOD) and Chemical Oxygen Demand (COD) are

used as wastewater quality indicators. There would be fees levied for trade effluent with BOD in

excess of 400mg/l to 4000mg/l. For BOD exceeding 4000mg/l, the trade effluent would have to

be treated to below 4000mg/l before discharging into public sewers. Effluent discharge has to be

treated to meet these limits before discharge.

10.6.3.3 Waste management & minimization

Waste management to be considered for this hydrogen plant would be the disposal of the

spent metal catalysts after catalyst changeout. Licensed general waste collectors will be

employed for this task. It is an offence for any person or company to collect or transport waste as

a business without a valid General Waste Collector License. The spent catalyst has to be

carefully handled, especially because catalyst fines might get into the air and cause health

problems to the people.

10.6.3.4 Energy efficiency

The National Energy Efficiency Committee (NEEC) is a committee with 3P (People,

Private, and Public Sector) representation. It seeks to integrate the promotion of energy

efficiency and the use of clean energy sources with the reduction of emissions of air pollutants

and carbon dioxide from the production of energy. The key objectives of the NEEC are as

follows:



• Promotion of energy conservation through efficient use of energy in the industrial,

building, and transportation sectors

• Promotion of the use of cleaner energy sources such as natural gas and renewable energy

sources

• Promotion of Singapore as a location for the pilot test-bedding of pioneering energy

technologies and as the hub for development and commercialization of clean energy

technologies

As can be seen from our mitigation measures, we are actively fulfilling the above mentioned

objectives of the NEEC.

10.6.4 Hydrogen Product Life Cycle Assessment

Fig 10-2: Hydrogen life cycle

Life cycle assessment (LCA) is defined as a systematic analytical method that helps

identify and evaluate the environmental impacts of a specific process or competing processes

[R12]. To quantitatively account for its impact on the environment, material and energy balances



are performed in a “cradle-to-grave” manner on the processes required to turn raw feedstock into

tangible products. Possible emissions, resource consumption and net energy consumption would

serve as primary indicators for the efficiency of the hydrogen life cycle. Finally, this LCA will be

used as the basis of comparison with other hydrogen generation methods to weigh the

environmental benefits and disadvantages of these various methods.

A study done by the National Renewable Energy Laboratory [R12], under the U.S.

Department of Energy, indicated that CO2 was emitted in the greatest amount, making up 99%

by weight of the total air emissions during steam methane reforming. This amount of CO2 also

contributed for 89.3% of the system’s global warming potential (GWP), defined as a

combination of CO2, CH4, and N2O emissions expressed as CO2-equivalence for a 100 year time

frame. Moreover, methane accounted for 10.6% of the GWP. Overall, the hydrogen plant itself

contributed 74.8% of the greenhouse gas emissions. Besides these gases, other hydrocarbons

(C2+), NOx, SOx, CO, particulates and benzene make up the remainder of the emissions. These

usually came about from natural gas production and distribution. Water was also consumed in

copious amounts in the hydrogen plant in the SMR, HTS and LTS reactors.

In terms of energy balance, it was determined that a major component of energy

consumption was found contained in the natural gas feedstock. On a life cycle basis, for every

MJ of fossil fuel consumed by the system, 0.66 MJ of hydrogen are produced on a LHV basis.

This figure has also included the energy used in the production, distribution of natural gas, and in

the generation of electricity to power the hydrogen plant itself [R12].

10.6.4.1 Ramifications of Hydrogen LCA

Hydrogen, an energy carrier, is perceived by engineers and scientists to be the energy

system for the 21st century. This is due to its abundance in the universe. However, H2 does not

exist naturally on Earth. It is mainly found on Earth as water and in organic compounds such as

methane, coal, petroleum and biomass.



At present, the majority of hydrogen is produced by the steam methane reforming method,

since it is still the most economical choice. However, this process relies heavily on liquid or gas

hydrocarbon fuels as the basic material for manufacture and as energy input for its production.

Therefore, the associated environmental impact is significant, as greenhouse gases such as CO2

and NOx would be emitted.

Moreover, another inherent problem in the product life cycle of hydrogen is related to the

lower energy density per unit volume of hydrogen. Hence, for hydrogen to be used as

transportation fuel, an energy-intensive liquefaction process is required. This not only incurs

additional cost, but also contributes to the emission of greenhouse gases, leading to increased

global warming.

Hence, although hydrogen is indeed a promising candidate as a future energy carrier,

with its reputation as a clean fuel with zero toxic emissions, the overall life cycle efficiency at

present is negative (-39.6%), i.e. the energy in the natural gas is greater than the energy content

of the hydrogen produced. It has also been mentioned above that for every MJ of fossil fuel

consumed by the system, 0.66 MJ of hydrogen are produced (LHV basis). 12 Therefore, unless

higher life cycle efficiencies are attained, the amount of resources, emissions, wastes and energy

consumption would remain a stumbling block towards its widespread implementation. Most

importantly, its adverse effect on the environment would still remain an issue to be resolved.

10.7 CONCLUSION

To conclude, it is an undeniable fact that safety issues should take precedence ahead of

economic considerations in the area of plant design and operation. The safe running of the

hydrogen plant will not only minimize human casualties and environmental harm, but can also

work hand-in-hand to meet economic demands being placed on the plant.



REFERENCES

[R1] : Daniel, A.C. & Louvar, J.F. (2002). Chemical Process Safety: Fundamentals with

Applications. 2nd Ed., Upper Saddle River, NJ: Prentice Hall.

[R2] : Mecklenburgh, J. C. (1985). Process Plant Layout, London: G. Godwin.

[R3] : Hydrogen Properties. Retrieved April 12, 2008, from U.S. Department of Energy Web

site: http://www1.eere.energy.gov/hydrogenandfuelcells/tech_validation/pdfs/fcm01r0.pdf

[R4] : Bausbacher, E. & Hunt, R. (1993). Process Plant Layout and Piping Design, New Jersey:

Prentice-Hall.

[R5] : Changi Water Reclamation Plant. Retrieved April 13, 2008, from CPG Corporation Web

site: http://www.cpgcorp.com.sg/portfolio/viewdetails.asp?Lang=EN&PCID=11&PDID=163

[R6] : Physical properties of gases, safety, MSDS, enthalpy, material compatibility, gas liquid

equilibrium. Retrieved April 14, 2008 from Air Liquide Web site:

http://encyclopedia.airliquide.com/encyclopedia.asp?CountryID=19&LanguageID=11

[R7] : Munn, R.E. (1979). Environmental Impact Assessment: Principles And Procedures. 2nd

Edition. New York: Wiley.

[R8] : Briffett, C. (1994). The Effectiveness of Environmental Impact Assessment in Southeast

Asia.

[R9] : Bass, R.E., Herson, A.I. and Bogdan, K.M. (1999) CEQA Deskbook: A Step-by-step Guide

on how to Comply with the California Environmental Quality Act. 2nd edn., Point Arena, CA:

Solano Press.

[R10] : OSH Answers: Legionnaire’s Disease. Retrieved on April 16, 2008 from Canadian

Centre for Occupational Health and Safety Web site:

http://www.ccohs.ca/oshanswers/diseases/legion.html

[R11] : Legionnaire’s Disease eTool : Source and Control – Cooling Towers, Evaporative

Condensers and Fluids Coolers. Retrieved on April 16, 2008 from U.S. Department of Labor

Occupational Safety & Health Administration Web site:

http://www.osha.gov/dts/osta/otm/legionnaires/cool_evap.html#Treatment



[R12] : Life Cycle Assessment Of Hydrogen Production Via Natural Gas Steam Reforming.

Retrieved on April 16, 2008 from National Renewable Energy Laboratory Web site:

http://www.nrel.gov/docs/fy01osti/27637.pdf

CN 4120: Design II Team 32 Instrumentation & Control


Chapter 11 : INSTRUMETNATION & CONTROL

11.1 INTRODUCTION

Instrumentation and control of the plant is critical to the plant’s operation and product

quality. It helps to manage the product quality through proper control of the plant. The objectives

of instrumentation and control are based on two major aspects [R1]:

Safety, Health and Environment (SHE) – A safe plant operation prolongs the life of the

expensive equipment and protects the health of the operators. Safe and smooth plant operation is

achieved with proper control that will detect abnormalities and effect the corrective actions to

maintain the process variable within the permissible limits. In addition, safe operation keeps the

surrounding environment in check by preventing unexpected harmful emissions to the

atmosphere.

Product quality – Through control and instrumentation, the plant would be able to respond to

changes in operating conditions quickly and effectively, thus ensuring minimal disturbances to

the plant. Therefore, the production scheme remains unperturbed and generates products of

constant quality and yield.

The protection scheme of the safety design is shown:

Fig 11-1: Typical layers of protection in a modern chemical plant



This chapter looks at the process control system of the steam methane reformer (SMR).

The control and instrumentation of this section is fully developed by our team and will be

explained in detail. The installation of critical alarms and automatic safety lock systems to

protect property and employees during emergencies are also explored. The controls are

employed bearing in mind that non-essential controls were reduced to minimize cost of the

process control system, without compromising on safety and quality of products.

11.2 PROCESS CONSIDERATION AND DESCRIPTION

Aims and objectives

The furnace and SMR reactor are two important units to the plant. The steam methane

reformer generates hydrogen while the furnace controls the SMR process. Controlling the steam

methane reforming process ensures constant yield and profitability of the plant. It also prevents

any possible runaway reactions. As the SMR process is endothermic, the heat duty is supplied by

the furnace and hence combustion is strictly controlled to maintain constant reaction temperature.

In addition, the high temperature and pressure associated with furnace operations creates more

necessity to impose proper control on the furnace. The lack of control of process variables within

the furnace may pose environmental and safety issues. The important process parameters related

to these two units are listed below:

Important process variables of Furnace

1. Temperature

2. Pressure

Important process variables of SMR reactor

1. Temperature

2. Steam methane ratio



11.3 PROCESS CONTROL METHODOLOGY

The following steps will be used in the formulation of our instrumentation and control design.

1. Identify the variables to be controlled, measured and manipulated

2. Select the control strategy and structure

3. State the controller settings

The general control strategies are feedback and feed-forward control. Feedback control

provides an easy control of variables without extensive knowledge of the process. Feed-forward

control provides a safer control as compared to feedback control because it allows corrective

action to be taken before the process variables go out of hand. The advantages and disadvantages

of feedback and feed-forward control are summarized in Table 11-1:

Advantages Disadvantages Feedback • Little knowledge is required

of the control process

• Disturbance need not be measured

• Corrective action will be taken regardless of the source and type of disturbances

• Poor control occurs if time lags are significant

• Closed-loop instability may occur

• Process upset takes place before corrective action is taken

Feed-forward • Corrective action is taken before process upset occurs

• “Perfect” control can be achieved.

• In depth knowledge of the process is required

• Ideal controllers may not be present to effect perfect control

Table 11-1: Summary of advantages and disadvantages of feedback and feed-forward control



11.4 SELECTION OF CONTROLLED, MANIPULATED AND MEASURED VARIABLE

A good control system can be achieved if the appropriate controlled and manipulated

variables are chosen. These are some guidelines we followed when selecting the variables [R1].

For controlled variable:

1. All variables that are not self-regulating must be controlled

2. Choose output variables that must be kept within equipment and operating constraints

3. Select output variables that represent a direct measure of product quality or that strongly

affect it

4. Choose output variables that interact with other controlled variables

5. Choose output variables that have favourable dynamic and static characteristics

For manipulated variable:

1. Select inputs that have large effects on controlled variables

2. Choose inputs that rapidly affect the controlled variables

3. The manipulated variables should affect the controlled variables directly rather than

indirectly

4. Avoid recycling of disturbances

For measured variable:

1. Reliable, accurate measurements are essential for good control

2. Select measurement points that have an adequate degree of sensitivity

3. Select measurement points that minimize time delays and time constraints



In our design, the major variables affecting the overall safety of the furnace are identified

as follows:

1. Pressure

Pressure control within the furnace is essential because it can compromise the

performance. Pressure drop within the reformer tubes has to be kept to a minimum for

favourable reaction conversion. Backflow within pipelines has to be avoided as well.

2. Temperature

Temperature control is essential to ensure that the reaction within the reformer tubes

proceed smoothly. It is also important to ensure that the materials of construction remain

intact by ensuring that the temperature within the furnace is kept below the creep

temperature of the materials. Flue gas temperature has to be controlled to prevent acid

gas condensation and to comply with governmental regulation.

3. Composition and Flow

The steam/methane ratio has to be kept above a minimum to prevent coking and high

pressure drop within the tubes. In addition, the air-to-fuel ratio into the furnace has to be

kept constant to ensure absolute combustion and maximum efficiency.

11.5 DETAILED CONTROL DESIGN FOR REFORMER FEED

11.5.1 Steam-to-Methane Ratio Control

One important control parameter in steam-methane reforming is the steam-to-methane

ratio. A low ratio is undesirable as it promotes the side reaction of coke formation on the catalyst,

which deactivates it and requires expensive replacement. Nonetheless, a high steam-to-methane

ratio will result in better conversion, but at the expense of elevated operating costs due to the

high cost associated with superheated steam. Hence, a compromise between methane conversion

and operating expense has to be made and in industries, this ratio is typically kept at 3:1.



Fig 11-2: Control scheme for steam/methane ratio control

To achieve this, a ratio control depicted in Fig 10-2 is employed to maintain the ratio

between steam and methane at 3:1 as stated in our design problem. The flow rate of the natural

gas stream is measured and transmitted by FT-101 to the ratio station FY-101. At the ratio

station, this signal is multiplied by an adjustable gain whose value is the desired ratio. The output

signal from the ratio station is then used as the set-point for flow controller FIC-102. This feed-

forward controller then adjusts the flow rate of the imported superheated steam by manipulating

the opening of the diaphragm valve FCV-102 using pneumatic signals.

In the preliminary design of this ratio control, it is assumed that molar flow rate is equal to the

volumetric flow rate which implies that possible pressure and temperature fluctuations in process

streams are not compensated.



11.5.2 Pressure Control Loop for Expander

A feedback control loop is used for all the expanders in the plant. The controlled variable

is the outlet pressure of the expanded vapour, while the manipulated variable is the inlet pressure

to the expander.

Fig 11-3: Control scheme for pressure control of expander discharge

The pressure transmitter (PT) would detect any deviations from the set point and send a

signal to the pressure controller (PIC) so as to adjust the valve which changes the inlet pressure

to the expander. This would then serve to bring the controlled variable back to its set point.

11.5.3 Temperature Control Loop to Preheat SMR Feed

The diagram below depicts a temperature control loop prior to the entry of the process

fluid into SMR. It is important to control the inlet temperature to SMR because it will affect the

methane conversion in the SMR. The inlet temperature to SMR can be controlled by varying the

flow of the SMR effluent through heat exchanger, HX-102. In this case, the temperature of the

preheated SMR feed is the controlled variable, while the flow of the SMR effluent is the

manipulated variable.

A cascade control loop is employed as follows:

1. The master controller is TIC-101, and the temperature of preheated SMR feed serves as

the set point for the slave controller, which is 539oC.

2. The two control loops are nested, with the secondary control loop (for the slave controller,

FIC-103) located inside the primary control loop (for the master controller, TIC-101)



Fig 11-4: Control scheme for temperature control loop to preheat SMR feed

The advantage of cascade control is to ensure a swift detection of any flow rate deviation

of the SMR effluent for which necessary adjustments will be made by FIC-103 before an upset in

the temperature of the preheated SMR feed can be effected. This provides a fast response to

deviations from set point values and a feed-forward control loop for the secondary control loop is

subsequently adopted for a faster response time.

11.5.4 Composition Analyzer for SMR Effluent

The composition analyzer A-004 can be used to determine reaction completion by

measuring the methane content of the effluent stream. The composition analyzer can be either an

infra-red or a chromatographic analyzer [R2] which measures a range of 0 to 10% methane in the

background of hydrogen and carbon monoxide. The advantage of using this composition

analyzer in that a desired furnace temperature profile can be arrived at by manipulating the

furnace temperature to achieve the desired degree of conversion.



11.6 DETAILED CONTROL DESIGN FOR SMR FURNACE

11.6.1 Air-to-Fuel Ratio Control

Operating without sufficient air can lead to fuel wastage due to inefficient combustion of

air and can even escalate to a hazard when the flammable products of incomplete combustion

ignite in the convection section.[3] However, excess air can reduce the efficiency of the furnace

due to large volumes of air heated to exit stack temperature without producing useful heat

transfer. From the above mentioned points, it can be seen that having an air-to-fuel ratio control

is crucial for the safe and efficient operation of furnace.

Fig 11-5: Control scheme for air-to-fuel ratio control

The air-to-fuel control scheme shown above is designed to fix the excess air at 15% to

ensure complete and stable combustion. In this control scheme, the disturbance variable is the

flow rate of the furnace fuel while the manipulated variable is the combustible air flow. The flow

rate of furnace fuel, which is a combination of PSA purge gas and natural gas, is measured and

transmitted by FT-106 to the ratio station FY-106. At the ratio station, this signal is multiplied by

an adjustable gain whose value is the desired ratio. The output signal from the ratio station will



act as the set-point for feed-forward controller FIC-107 to control the valve opening of

diaphragm valve FCV-107 to adjust the flow of combustion air to the burners of furnace so as to

maintain the excess air at 15%.

11.6.2 Temperature Control Loop to Regulate Effluent Exit Temperature

Natural gas is burned in the furnace to supply the heat duty required for the endothermic

conversion of methane to hydrogen. Therefore, in this control system, the manipulated variable is

the flow rate of natural gas, while the controlled variable is the temperature of the SMR effluent

being heated.

Fig 11-6: Control scheme for temperature loop to regulate effluent exit temperature

A cascade control loop is employed as follows:

1. The master controller is TIC-105, and the temperature of SMR effluent serves as the set point

for the slave controller

2. The two control loops are nested, with the secondary control loop (for the slave controller,

FIC-105) located inside the primary control loop (for master controller, TIC-105)



This cascade control will ensure that any deviations in the flow rate of natural gas would

be detected and FIC-105 would make adjustments to the flow rate of natural gas even before it

could cause an upset to the SMR effluent temperature. This is the advantage of cascade controls,

whereby a second measured variable is located close to the potential disturbance and its

associated controller reacts quickly, where it offers very fast response time to deviations from set

point values. To improve the response time further, a feed-forward control loop for the secondary

control loop was employed.

11.6.3 Pressure Control Loop to Regulate Furnace Draft

Fig 11-6: Control scheme for pressure control loop to regulate furnace draft

A pressure control loop is designed to maintain a small negative pressure at the top of the

radiant section just before the convection section. This is because if the pressure is positive in the

radiant section, it will cause the hot flue gas to leak outward and damage the steel structure of

furnaces, thus shortening the lifespan of furnace.

The pressure can be maintained at negative pressure by adjusting the opening of the stack

damper using a feedback control. When the pressure before the convection-section inlet deviates

from the set-point value, the pressure transmitter PT-001 will send a signal to pressure controller

PIC-001 in the feedback control loop which will then adjust the opening of the stack damper to



either increase or decrease the draft. In general, when the stack damper closes, the draft

decreases and vice versa. A position transmitter ZT-001 will measure the opening of the stack

damper and reflect it via the indicator ZI-001.

11.6.4 Flue Gas Exit Temperature Control

Flue gas exit temperature control is necessary for two reasons: 1) to prevent corrosion

attack caused by acid gas condensation and 2) to conform to governmental regulation for flue gas

exit temperature. The temperature control at the stack exit is achieved by a cascade control

configuration which consists of a primary control loop utilizing TT-104 and TIC-104 and a

secondary control loop that controls the flow of condensate via FT-104 and FIC-104.

The exit flue gas temperature will be measured by the temperature transmitter TT-104

and will be used by the master controller TIC-104 to establish a set point for the secondary loop

controller, FIC-104.

Fig 11-7: Control scheme for temperature control loop for flue gas exit



The reason why cascade control is used here instead of conventional feedback control is

because if feedback control is used and there is variation in the flow of the condensate steam

which cause a change in flue gas temperature, this change can only be effected after the

temperature controller takes corrective action to adjust the condensate flow. In our case where

cascade control is used, the flow controller FIC-104 will respond very fast to hold the condensate

flow at its set point without causing any disturbances to the flue gas temperature.

11.6.5 Tube Metal Temperature Indicator

Reformer tubes are designed to operate at a particular pressure and temperature.

Conditions such as flame impingement, poor radiant-heat distribution, interior tube deposits

(coke) can cause localized overheating of tube known as hot-spots, which could result in high

temperature creep. Hence, it is mandatory to have a temperature monitoring device to monitor

furnace tube skin temperature to prevent the plastic deformation of tubes.

Fig 11-8: Tube metal temperature indicators

To achieve this, thermocouples can be welded onto reformer tubes. In addition,

temperature high alarms should also be incorporated together with the thermocouples to alert

operators of high tube skin temperature. This practice is adopted in our design of the steam

methane reformer as a safeguard against tube failure. A common practice in industry requires the



installation of thermocouples on 15% of the reformer tubes but due to space constraints on P&ID,

only 4 thermocouples are reflected.

11.6.6 Analyzers for Furnace Control

Analytical devices such as oxygen analyzers and combustible detectors are used to

control the operation of the furnace. The oxygen sensor A-002, which is a solid state heated

zirconium oxide probe that is inserted directly in the stack, can be used to measure the oxygen

content in the flue gas in order to compute the excess air that is being delivered. As such, the

efficiency of furnace combustion can be maximized by throttling the air entering the furnace to

maintain the oxygen content at a desirable value. The sensor can also be tied to a low oxygen

alarm to warn operators if a hazardous furnace atmosphere is developing.

Furthermore, a combustible analyzer A-003, which uses an infra-red beam directed

across the stack to measure the amount of carbon monoxide content in the flue gas, is also

utilized for furnace control. The combustible analyzer usually works in tandem with an oxygen

analyzer to provide a correlation between excess air and unburned fuel going up the flue stream.

For this reason, they can serve as a form of backup to check for the effectiveness of the

air-to-fuel ratio control loop. Although combustible analyzers can sometimes be used to

automatically adjust the air/fuel ratio, this control scheme was not adopted in our control of

furnace due to reliability issue of the combustible analyzer.

11.7 Safety Devices

11.7.1 Pressure Relief Valves

Relief valves are required to provide an outlet for over-pressurised fluids to prevent

rupture of pipelines so as to protect operators from possible hazards of over-pressurizing

equipment. In addition, relief valves also prevent damage to adjoining property; reduce insurance

premiums and chemical losses during pressure upset. With these in mind, it is in fact stipulated

by governmental regulations that installations of pressure relief systems are compulsory.



In this preliminary piping and instrumentation diagram (P&ID) design of the SMR

reactor/furnace, the pressure relief devices are designed to be installed at every point identified

as potentially hazardous. Since most of the pipelines are in gas service, safety valves are chosen

for pressure relief. Safety valves pop open to release the excess pressure when the operating

pressure exceeds the set pressure. The only exception is the pipeline with steam in service.

Hence a safety relief valve, which functions as relief valve for liquid and safety valve for steam,

may be required [R4].

11.7.2 Process Alarms

The function of a process alarm is to warn operators of impending dangers when process

parameters such as temperature, pressure, flow or level exceed or fall below the permissible

limits. In general, alarms can be categorised as follows [R5]:

Type Description Function

Type 1 Alarm Equipment status

alarm

Indication to whether an equipment is switched

on or off

Type 2 Alarm Abnormal

measurement alarm

When activated, it acts as an indication that the

reading taken by the sensor is outside

acceptable limit

Type 3 Alarm An alarm without

an adjoining sensor

Alarm that is directly triggered by process

instead of sensor signal when the process

parameter is out of specification. Knowledge of

actual process value is not required

Type 4 Alarm An alarm with an

adjoining sensor

Serve as a backup to the regular sensor in the

event that it fails

Type 5 Alarm

Automatic

shutdown or start-

up system

Typical type of alarm that is widely used

Table 11-2: Summary of different categories of alarm



Upon activation of the alarm, an annunciator, either in the form of visual displays

(flashing light on control panels) or audible sounds (horns or bells), will be triggered. Unless

acknowledged by plant operators, the alarm will remain in activation. If the abnormal situation

that sets off the alarm is deemed to be potentially dangerous, an automated corrective action will

be initiated by the safety interlock system to shut down the affected unit.

Excessive number of unnecessary alarms should be avoided because frequent “nuisance

alarms” make plant operators less responsive to crucial alarms and may obscure the root cause of

the abnormal situation in the presence of many unimportant alarms. Hence, alarms should only

be installed at locations deemed absolutely necessary. The alarms in the P&ID and their

functions are summarized in the following table:

Alarm Tag No. Description for Alarm Cause THA -104 Flue Gas High Temperature

THA-201, THA-202, THA-203, THA-204 SMR Furnace High Tube Skin Temperature

ALA-002 Low Flue Gas Oxygen Content

PHA -001 SMR Furnace High Pressure

FLA -201 Fuel Gas Low Flow

FLA -202 Combustion Air Low Flow

FLA -203 SMR Feed Low Flow

FLA -204 SMR Effluent Low Flow

Table 11-3: Summary of alarms in our P&ID and their causes

11.7.3 Safety Interlocks or Emergency Shutdown System (SIS or ESD)

The control loops that are designed in the P&ID form the basic process control system

(BPCS) which acts as primary protection against deviations in process parameters. During

normal operations, BPCS can provide acceptable control but this may not be the case during

abnormal or emergency situations. For this reason, implementation of SIS or ESD is required to

serve as a backup especially when BPSC components malfunction or when there is utility failure.

As such, considerations must be made for SIS and ESD to function independently of the BCPS.



In view of the drastic measure such as the complete shutdown of process unit, the SIS

and ESD can only be considered as the last resort to protect the equipment and prevent injury to

personnel and used only when critical process variables go beyond the specified allowable

operating limit. Although SIS is essential for safe plant operation, unnecessary plant shutdown

should be precluded as it can reduce throughput due to downtime and may cause products to go

off-specification during subsequent plant start-up.

11.7.3.1 Implementation of SIS or ESD for the protection of nickel catalyst

Fig 11-9: SIS control scheme for protection of nickel catalyst

As mentioned earlier, a low steam/methane ratio is undesirable because it can deactivate

the nickel catalyst due to coke formation. To prevent this situation from arising, a SIS should be

in place to tackle any abnormalities in the ratio. The SIS to be implemented is illustrated in the

above control scheme. When the steam/methane ratio falls below the limit of approximately 3:1,

a flow switch FSL-001 will trigger off an alarm. If the ratio continues to fall below the critical

limit (approximately 2.7:1), the natural gas stream will be shut off via FSL-002, which trips a

solenoid that control the transducer I/P-101 and causes FCV-101 to be closed.



In designing such a SIS, several points must be taken into consideration. Firstly, FCV-

101 must be a quick-closing valve (4 to 5 seconds for full closure) so that the flow of natural gas

can be almost instantly stopped when there is a large deviation in reforming steam, thus

protecting the reformer catalyst. FCV must also be a single-seated tight shut-off valve, to prevent

leakage during the time while the electric-operated shutoff valve is in the process of closing.

11.8 Additional Considerations in Process Control

11.8.1 Redundancy of Air Blowers and Expanders

The reliability of plant equipment is critical to successful plant operation. Hence, spare

air blowers and expanders are provided in parallel to the commissioned equipment and are

placed on hot standby to ensure the continuous operation of the hydrogen plant whenever any of

the equipment undergo mechanical failure or are sent offline for maintenance purposes, which is

a common phenomenon during normal operation.

Fig 11-10: Redundancies used in process control



11.8.2 Isolation Valves and Bypass

Isolation valves, most commonly gate valves, are installed upstream and downstream of

process equipment such as heat exchangers, air pre-heater, control valves and air blower to

ensure the operability of the plant whenever there is maintenance or repair of the equipment.

When isolated, a bypass as illustrated must be present to divert the process flow to the other side

of the process equipment.

Fig 11-11: Bypass of pipelines

11.9 REFERENCES

[R1]: Seborg D. S., Edgar T. F., Mellichamp D. A. Process Dynamics and Control, John Wiley

& Sons (2004)

[R2]: Liptak B. Instrument Engineer’s Handbook: Process Control and Optimization, CRC Press

(2005)

[R3]: Lieberman N. P., Lieberman E. T., Working Guide to Process Equipment. 2nd Edition Mc-

Graw-Hill

[R4]: Crowl D.A., Louvar J.F. Chemical Process Safety: Fundamentals with Applications. 2nd

Edition, Prentice Hall PTR (2002)

[R5]: Connell, B., “Process Instrumentation Applications Manual”, Mc-Graw-Hill, New York (1996)

B/P

NATURAL GASFROM PIPELINE

IMPORTED STEAM

TT101

42 BAR HP STEAMFOR EXPORT

CONDENSATE

TE102

TE103

I/P102

I/P101

SMR EFFLUENT TOHX-102

TE105

TT105

I/P105

FT101

FT102

FT105

FT106

I/P107

FT107

PSA

PU

RG

E G

AS

( CO

2 re

mov

ed)

MAKEUP NATURAL GASTO FURNACE

COMBUSTION AIR

TE104

TT104

I/P104

FT104

ZT001

PT001

I/P001

F-101

AP-101 AP-102

HX-101

HX-102

102

FIC

101

FY

101

FIC

101

TIC

001

ZI001 opacity

A

002 O2

A

003 CO

A

105

TIC

106

FY

107

FIC

105

FIC

104

TIC

001

PIC

102

TI

103

TI

AIR

FR

OM

IN

TAK

E FI

LTER

104

FICB-101

B-102

TE201201

TI

201

THA TE202

202

TI

202

THA

TE203

203

TI

203

THA

TE204

204

TI

204

THA

101

PI

102

PI

E-101

E-102

E-103

E-104

E-105

E-106

AT004

004 Methane

A

SMR EFFLUENT

103

FIC

I/P103 FT

103

201

FI

201

FLA

001

PHA

FT202 202

FI

202

FLA

FT203

203

FI

203

FLA

FT204

204

FI

204

FLA

104

THA

PSV-001

PSV-002

PSV-003

PSV-004

PSV-005

PSV-006

PSV-007

PSV-008

PSV-009

PSV-010

PSV-011

PSV-012

PT

PIC

See Detail A

Detail A

Applicable to all expanders

002 O2

ALA

team 32 - overall team report

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