Thesis

Introduction to the Sweetening of Natural Gas with Emphasis on Sulfur Recovery (Sulfur Recovery: 80 tons per day)

Project Advisor:

Prof. Dr. Shahid Naveed

Project Co advisor:

Madam Masooma Sundus

Project Team

Name Registration No.

Imran Shabbir 2005-Chem-79

Omer Farooqi 2005-Chem-97

Jahanzaib Ali Bugti 2005-Chem-95

Ali Shan Malik 2005-Chem-41

Osman Shahid 2005-Chem-65

UNIVERSITY OF ENGINEERING AND

TECHNOLOGY LAHORE PAKISTAN

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Approval Certificate

INTRODUCTION TO THE SWEETENING OF NATURAL GAS

WITH EMPHASIS ON SULFUR RECOVERY

This major project report has been completed and submitted to the Department of Chemical

Engineering, University of Engineering and Technology Lahore in partial fulfillment of the

requirement for the B.Sc Chemical Engineering degree

Project Team:

Imran Shabbir

Muhammad Omer Farooqi

Jahanzaib Ali Bugti

Ali Shan Malik

Osman Shahid

Approved by:

Prof. Dr. Shahid Naveed Prof. Dr. A. R. Saleemi

(Project Advisor) (Chairman)

____________________ ___________________

External Examiner

____________________ ___________________

In the name of

Allah The most Merciful and

Compassionate, The most Gracious and Beneficent Whose help and

guidance we always solicit at every step, in each moment of our lives

DEDICATION

Our parents, whose

blessing brought us at

this stage and who

trample their

inclination & longings

for uploading our

studies

Acknowledgments Thanks to The Almighty ALLAH, “Who taught us with pen and

told what we did not know” and guided us by the Holy Prophet Hazrat

Mohammad (Peace be upon Him) after whom no further guidance is

needed.

We are indebted to our chairman, Prof. Dr. A. R. Saleemi who

provided us his knowledge and facilities to complete this project.

We acknowledge our indebtedness to our beloved project adviser

Prof. Dr. Shahid Naveed and our project co adviser Madam Masuma

Sundus for their timely guidance, encouragement, sympathetic attitude

and professional assistance, without which this project would not have

been completed.

A special thank you goes to Engr. Sir Mohsin Kazmi, Engr. Sir

Faheem and Engr. Sir Qazi Zaka ur Rehman for being so kind and

helping to us. Indeed without their guidance it was not an easy job to

complete this project.

A general debt of gratitude is due to all the teachers of the

Chemical Engineering Department, UET Lahore for their kind help.

There is a deep contribution from our teachers to whatever we have

achieved and whatever we intend to achieve in our lives.

We are also thankful to the non-teaching staff of the department

for their intellectual and moral support.

We extend special thanks to our sweet parents for their unlimited

love, kindness and support throughout our studies, and who pray for our

success and bright future deeply.

AUTHORS

TABLE OF CONTENTS

Abstract І

Preface І І

Problem Statement (1)

Chapter 1 (2-20)

INTRODUCTION TO NATURAL GAS PROCESSING

1.1 Exploration of Natural Gas

1.2 Processing Natural Gas

1.3 Sweetening

1.3.1 Reasons of Removing H2S and CO2

1.3.2 Amine Solutions used in Sweetening

1.3.3 The Girdler Process

1.4 About Sulfur

1.4.1 Properties of Sulfur

1.4.2 Processing of Sulfur

1.4.3 Uses of Sulfur

1.4.4 Products of Sulfur

1.5 Sulfur Recovery Methods

1.5.1 Medium (0.20 to 25.0 LTPD)

1.5.2 Large (greater than 25.0 LTPD)

1.5.3 Explanation of Various Processes

1.5.3.1 Sulfa Treat Direct Oxidation Process

1.5.3.2 The Claus Process

1.5.3.3 Recycle Selectox Process

1.5.3.4 Selective Oxidation Process

1.5.3.5 Cold Bed Adsorption Process

1.5.3.6 Thermal Cracking of H2S

1.6 The Claus Process

1.6.1 History

1.6.2 Description

1.6.3 Simplified Process Description

1.6.4 Process Improvements

1.6.5 Claus Process Auxiliaries

1.6.5.1 Blow Down System

1.6.5.2 Fuel System

Chapter 2 (21-30)

MAJOR EQUIPMENTS USED IN CLAUS PROCESS & THEIR

IMPROVEMENT CONSIDERATIONS

2.1 Introduction

2.2 Reaction Furnace (F-100)

2.3 Waste Heat Boiler (B-100)

2.4 Sulfur Condensers (E-100, E-102, E-104, E-106)

2.5 Heaters (E-101, E-103, E-105)

2.5.1 Direct Reheat Methods

2.5.2 Hot Gas Bypass

2.5.3 Acid Gas Fired Line Burner

2.6 Catalytic Reactors (R-100, R-101, R-102)

2.7 Sulfur Pits

Chapter 3 (31-45)

MATERIAL BALANCE OF THE SULFUR RECOVERY UNIT (SRU)

3.1 Introduction

3.2 Overall Material Balance

3.3 Material Balance across Furnace F-100

3.4 Material Balance across Condenser E-100

3.5 Material Balance across Reactor R-100






3.11 Final Calculations

Chapter 4 (46-57)

ENERGY BALANCE OF THE SULFUR RECOVERY UNIT (SRU)

4.1 Introduction

4.2 Overall Energy Balance

4.3 Energy Balance across Furnace F-100

4.4 Energy Balance across Boiler B-100

4.5 Energy Balance across Condenser E-100

4.6 Energy Balance across Heater E-101

4.7 Energy Balance across Reactor R-100








Chapter 5 (58-93)

EQUIPMENTS DESIGN

5.1 Design of Reaction Furnace (F-100)

5.2 Design of Waste Heat Boiler (B-100)

5.3 Design of Reactors (R-100, R-101, R-102)

5.4 Design of Condenser (E-106)

5.5 Design of Process Stream Heater (E-105)

Chapter 6 (94-97)

PROCESS INSTRUMENTATION & CONTROL

6.1 Introduction

6.2 General Discussion on the Instrumentation of the Sulfur Recovery Unit (SRU)

6.2.1 Feed Flow Measurement and Control

6.2.2 Combustion Air Control

6.2.3 Main Burner and Reaction Furnace

6.2.4 Waste Heat Boiler

6.2.5 Sulfur Condensers

6.2.6 Heaters

6.2.7 Catalytic Reactors

6.2.8 Shutdown System

6.3 Instrumentation for Condensers

Chapter 7 (98-101)

MECHANICAL DESIGN OF THE PROCESS STREAM HEATERS

7.1 Introduction

7.2 Waste Heat Boiler

7.3 The Claus Reactors

7.4 Sulfur Condensers

7.5 Sulfur Pits

Chapter 8 (102-104)

HAZOP STUDY OF THE SULFUR RECOVERY UNIT (SRU)

8.1 Introduction

8.2 General Safety Rules

8.3 Building and Process Equipment Safety

8.3.1 Lights

8.3.2 Electrical and Mechanical Hazards

8.3.3 Chemical Hazards

8.3.4 Fire Prevention and Control

8.3.5 Personnel Safety

8.4 Claus Process

8.4.1 Special Hazards and Precautions

Chapter 9 (105-112)

COST ESTIMATION OF THE SULFUR RECOVERY UNIT (SRU)

9.1 Introduction

9.2 Fixed and Working Capital

9.3 Total Production Cost

9.3.1 Manufacturing Cost

9.3.2 General Expenses

9.4 Equipment Cost

References

I

ABSTRACT

This project has to design Sulfur Recovery Unit (SRU). There are many

processes for the recovery of sulfur from natural gas but we selected the Claus

process, because the design of the process is economically most favorable. The

economics of the plant also make balance with the efficiency and is most suited

to Pakistan’s wells of oil and gas.

The process selected for this purpose is the Claus process and the unit is

designed to produce 80 tons of elemental sulfur per day.

This report includes introduction to natural gas exploration, Dakhni gas

processing plant review, production and processing, various processes

employed for the sulfur recovery from the natural gas, the details of the Claus

process, material and energy balances across the sulfur recovery unit (SRU),

individual equipments design, instrumentation and control, piping, cost

estimation, selection process of the construction material and lining of

refractory, and the safety of the sulfur recovery unit (SRU).

In summary, the focus on the future improvements in the Claus process

makes this project distinctive and particularly relevant for educating present or

perspective engineers. We have worked hard to complete this project that is

stimulating for engineers to read. We also strived to develop the design of the

sulfur recovery unit (SRU) that will capture engineer’s attention, is

pedagogically sound and well integrated with project material, and is easy for

the engineers to use and adapt.

We welcome any comments or suggestions. Please feel free to contact via

e-mail at: [email protected] and

[email protected], furthermore a soft copy can be obtained on

request at the said e-mail addresses.

II

PREFACE

The aim behind this project is to design the sulfur recovery unit (from

natural gas). The capacity of the proposed plant is 80 tons per day.

Generally, the natural gas obtained from the reservoirs, contains many

impurities including hydrogen disulfide (H2S), the presence of which makes the

gas toxic. To make the use of this gas environmentally acceptable, the gas is

passed through a number of purifying stages. One of these stages is that of

sulfur recovery unit (SRU).

There are many different processes used for the recovery of sulfur from

natural gas. We selected the Claus process, as it is the most economical process

especially for the large amounts to process like we had to. This process mainly

comprises two reactions; first, one by third of the hydrogen disulfide present in

the feed is converted into sulfur dioxide by burning in the furnace and second,

the remaining hydrogen disulfide reacts with the produced sulfur dioxide to

give elemental sulfur. First reaction occurs in a furnace while the second

reaction takes place in a series of reactors. Sulfur produced in the reactors is

then condensed in the condenser. The pipelines throughout are insulated so that

sulfur may not freeze inside the pipes.

Sulfur obtained by this process is used commercially as a hardening

agent in the manufacture of rubber products, such as tires. The most important

use of sulfur is in the manufacture of sulfur compounds, such as sulfuric acid,

sulfites, sulfates, and sulfur dioxide. Medicinally, it has assumed importance

because of its widespread use in sulfa drugs and in many skin ointments. Sulfur

is also employed in the production of matches, wood pulp, carbon disulfide,

insecticides, bleaching agents, vulcanized rubber etc.

1

Problem Statement

This project report had been assigned to us as the partial fulfillment for the requirement of the B.Sc Chemical Engineering degree. The problem statement is: “INTRODUCTION TO THE SWEETENING OF NATURAL GAS WITH EMPHASIS ON SULFUR RECOVERY”. The proposed plant capacity is selected to be 80 tons per day which matches with the prevailing extended market needs and to meet the industrial demands. The inspiring facility for this project is the Oil and Gas Development Corporation Limited (OGDCL), Dakhni. This facility has a current production of 65 tons per day of elemental rhombic sulfur but is interested in extension of the production plants to produce 80 tons per day which is the very problem assigned to us in this project.

The natural gas obtained from wells contains toxic hydrogen sulfide gas which must be removed in order to make the use of natural gas safe and friendly. Sweetening is done to remove hydrogen sulfide gas and then the famous Claus process is employed to recover elemental rhombic sulfur from the hydrogen sulfide gas stream which is a valuable market product having its use, in the production of many daily life useful products, as a raw material.

Fig: Flow diagram for the Claus process.

The project team is guided and motivated by respected Dr. Shahid Naveed as the project advisor. All of the material and data being presented in this report is taken from authentic literature and timely references have been provided to guide the reader and at the same time prevent ourselves of getting divert from the main essence of report writing.

Chapter 1 INTRODUCTION AND LITERATURE REVIEW

1.1 Exploration of Natural Gas

he practice of locating natural gas and petroleum deposits has been transformed dramatically in the last 15 years with the advent of extremely advanced, ingenious technology. In the early days of the industry, the only way of locating underground petroleum and natural gas deposits was to search for surface evidence of these underground formations. Those searching for natural gas

deposits were forced to scour the earth, looking for seepages of oil or gas emitted from underground before they had any clue that there were deposits underneath. However, because such a low proportion of petroleum and natural gas deposits actually seep to the surface, this made for a very inefficient and difficult exploration process. As the demand for fossil fuel energy has increased dramatically over the past years, so has the necessity for more accurate methods of locating these deposits.

1.2 Processing Natural Gas

A Natural Gas Processing Plant Natural gas, as it is used by consumers, is much different from the natural gas that is brought from underground up to the wellhead. Although the processing of natural gas is in many respects less complicated than the processing and refining of crude oil, it is equally as necessary before its use by end users. The natural gas used by consumers is composed almost entirely of methane. However, natural gas found at the wellhead, although still composed primarily of methane, is by no means as pure. Raw natural gas comes from three types of wells: oil wells, gas wells, and condensate wells. Natural gas that comes from oil wells is typically termed as “associated gas”. This gas can exist separate from oil in the formation (free gas), or dissolved in the crude oil (dissolved gas). Natural gas from gas and condensate wells, in which there is little or no crude oil, is termed ‘non-associated gas’. Gas wells typically produce raw natural gas by itself, while condensate wells produce free natural gas along with a semi-liquid hydrocarbon condensate. Whatever the source of the natural gas, once separated from crude oil (if present) it commonly exists in mixtures with other hydrocarbons; principally ethane, propane, butane, and pentanes. In addition, raw natural gas contains water vapor, hydrogen sulfide (H2S), carbon dioxide, helium, nitrogen, and other compounds.

Natural gas processing consists of separating all of the various hydrocarbons and thuds from the pure natural gas, to produce what is known as ‘pipeline quality’ dry natural gas. Major transportation pipelines usually impose restrictions on the make-up of the

T

Chapter 1 Introduction to Natural Gas Processing

3

natural gas that is allowed into the pipeline. That means that before the natural gas can be transported it must be purified. While the ethane, propane, butane, and pentanes must be removed from natural gas, this does not mean that they are all ‘waste products’.

In addition to processing done at the wellhead and at centralized processing plants, some final processing is also sometimes accomplished at ‘straddle extraction plants’. These plants are located on major pipeline systems. Although the natural gas that arrives at these straddle extraction plants is already of pipeline quality, in certain instances there still exist small quantities of NGLs, which are extracted at the straddle plants.

The actual practice of processing natural gas to pipeline dry gas quality levels can be quite complex but usually involves four main processes to remove the various impurities:

Oil and Condensate Removal

Water Removal

Separation of Natural Gas Liquids

Sulfur and Carbon Dioxide Removal

Fig 1.1: Diagram of a typical gas processing plant.


4

1.3 Sweetening

Amine gas treating, also known as gas sweetening and acid gas removal, refers to a group of processes that use aqueous solutions of various alkanolamines (commonly referred to simply as amines) to remove hydrogen sulfide (H2S) and carbon dioxide (CO2) from gases. It is a common unit process used in refineries, petrochemical plants, natural gas processing plants and other industries. Processes within oil refineries or natural gas processing plants that remove hydrogen sulfide and/or mercaptans are commonly referred to as sweetening processes because they result in products which no longer have the sour, foul odors of mercaptans and hydrogen sulfide.

1.3.1 Reasons of Removing H2S and CO2

Carbon dioxide, hydrogen sulfide, and other contaminants are often found in natural gas streams. CO2 when combined with water creates carbonic acid which is corrosive. CO2 also reduces the BTU value of gas and in concentrations of more that 2% or 3% the gas is unmarketable. H2S is an extremely toxic gas that is also tremendously corrosive to equipment. Amine sweetening processes remove these contaminants so that the gas is marketable and suitable for transportation. The recovered hydrogen sulfide gas stream may be:

Vented to atmosphere.

Flared in waste gas flares or modern smokeless flares.

Incinerated for sulfur removal.

Utilized for the production of elemental sulfur or sulfuric acid.

If the recovered H2S gas stream is not to be utilized as a feedstock for commercial applications, the gas is usually passed to a tail gas incinerator in which the H2S is oxidized to SO2 and is then passed to the atmosphere out a stack.

1.3.2 Amine Solutions Used in Sweetening

Amine has a natural affinity for both CO2 and H2S allowing this to be a very efficient and effective removal process. There are many different amines used in gas treating:

Monoethanolamine (MEA) -Used in low pressure natural gas treatment applications requiring stringent outlet gas specifications.

Diethanolamine (DEA) -Used in medium to high pressure treating and does not require reclaiming as do MEA and DGA systems.

Methyldiethanolamine (MDEA) -Has a higher affinity for H2S than CO2 which allows some CO2 "slip" while retaining H2S removal capabilities.

Diisopropylamine (DIPA) Aminoethoxyethanol / diglycolamine (DGA) Formulated special solvents.

http://en.wikipedia.org/wiki/Hydrogen_sulfide

http://en.wikipedia.org/wiki/Carbon_dioxide

http://en.wikipedia.org/wiki/Unit_process

http://en.wikipedia.org/wiki/Oil_refinery

http://en.wikipedia.org/wiki/Petrochemical

http://en.wikipedia.org/wiki/Natural_gas_processing



http://en.wikipedia.org/wiki/Mercaptans

http://en.wikipedia.org/wiki/Ethanolamine

http://en.wikipedia.org/wiki/Diethanolamine

http://en.wikipedia.org/wiki/MDEA

http://en.wikipedia.org/wiki/Diisopropylamine

http://en.wikipedia.org/w/index.php?title=Aminoethoxyethanol&action=edit&redlink=1


5

However, the most commonly used amines in industrial plants are the alkanolamines MEA, DEA, and MDEA. Amines are also used in many oil refineries to remove sour gases from liquid hydrocarbons such as liquefied petroleum gas (LPG).

1.3.3 The Girdler Process

Natural gas is considered "sour" if hydrogen sulfide (H2S) is present in amounts greater than 5.7 milligrams per normal cubic meters (mg/Nm3) or 0.25 grains per 100 standard cubic feet [gr/100 scf]). The H2S must be removed (called "sweetening" the gas) before the gas can be utilized. If H2S is present, the gas is usually sweetened by absorption of the H2S in an amine solution, also known as the Girdler process.

Other methods, such as carbonate process, solid bed absorbent and physical absorption, are employed in the other sweetening plants.

The main reaction of the Girdler process is as follows:

2RNH2 + H2S (RNH3)2S

Where:

R = mono, di, or tri-ethanol

N = nitrogen

H = hydrogen

S = sulfur

Fig 1.2: A typical amine gas sweetening plant.

http://en.wikipedia.org/wiki/Sour_gas

http://en.wikipedia.org/wiki/Liquified_petroleum_gas


6

1.4 About Sulfur

Sulfur is a non-metallic element that occurs in both combined and free states and is distributed widely over the earth’s surface. It is tasteless, odorless, insoluble in water, and often occurs in yellow crystals or masses. It is one of the most abundant elements found in a pure crystalline form. The word sulfur is Latin for “burning stone”, and was used almost interchangeably with the term for fire. Because of its combustibility, sulfur was used for a variety of purposes at least 4,000 years ago.

Although it is plentiful on a world scale, native sulfur is usually found in relatively minute quantities. The greatest quantity of naturally occurring sulfur by far is combined with other elements, most notably the sulfides of copper, iron, lead, and zinc, and the sulfates of barium, calcium (commonly known as gypsum), magnesium, and sodium.

In the late 1 800s the Frasch process - a mining technique that recovers from 75% to 92% of a salt dome’s recoverable sulfur - became operational. Stockpiles today account for more than 50% of the. US, Canada, Japan, France, Poland, and Mexico are major sulfur suppliers.

Secondary sources of sulfur today are the sulfur dioxide (SO2) obtained from industrial mineral, wastes, and flue gasses, and the hydrogen sulfide (H2S) found in “sour” natural gas, petroleum refinery products, and coke-oven gasses. Once considered unwelcome byproducts of industrial processes, these sources of sulfur have the advantage of being nearly inexhaustible.

It was stated that 80% to 85% sulfur production in the year 2000 was recovered sulfur produced from hydrogen sulfide (H2S).

1.4.1 Properties of Sulfur

Chemical Name: Sulfur

Family Name: Element - Sulfur

Chemical Formula: S8

Physical State: Solid

Appearance: Yellow colored lumps, crystals, powder, or formed shape

Odor: Odorless, or faint odor of rotten eggs if not 100% pure

Purity: 90% - 100%

Molecular Weight: 256.50


7

Vapor Density: (Air = 1): 1.1

Vapor Pressure: 0 mmHg at 280 oF

Solubility in Water: Insoluble

Specific Gravity: 2.07 at 70 oF

Boiling Point: 832 oF (444 oC)

Freezing/Melting Point: 230-246 oF (110-119 oC)

Bulk Density: Lumps 75-1 15 lbs./ft3 Powder 3 3-80 lbs./ft3

Flashpoint: 405 oF (207.2 oC)

Flammable Limits: LEL: 3.3 UEL: 46.0

Auto-ignition Temperature: 478-511 oF (248-266 oC)

Sulfur is an odorless, tasteless, light yellow solid. It is a reactive element that given favorable circumstances combines with all other elements except gases, gold, and platinum. Sulfur appears in a number of different allotropic modifications: rhombic, monoclinic, polymeric, and others. The rhombic structure is the most commonly found sulfur form. Each allotropic form differs in solubility, specific gravity, crystalline, crystalline arrangement, and other physical constants. These various allotropes also can exist together in equilibrium in definite proportions, depending on temperature and pressure.

1.4.2 Processing of Sulfur

Sulfur processing is accomplished in plants using four manufacturing methods, producing sulfurs described as: Milled sulfurs, Formed sulfurs, Emulsified sulfur, and Precipitated sulfur.

Milled Sulfurs

These products are produced using Raymond roller mills to grind to specific particle ranges. Additives such as dispersants, flow aids, and dust suppressants may be added to enhance product performance.

Formed Sulfurs

These products are produced by molding, drum flaking, or rotoforining, and then sized to meet specific needs. Additives may be used for degradability.


8

Emulsified Sulfur

These products are manufactured using homogenizing technology to create a water based suspension.

Precipitated Sulfur

Flowers of sulfur are distilled sulfurs of exceptional purity obtained by sublimation of sulfur vapor into particulate form in an inert atmosphere.

1.4.3 Uses of Sulfur

Sulfur is an element used for everything from adhesives to matches. Its most common use is as a hardening agent in the manufacture of rubber products, such as tires. The most important use of sulfur is in the manufacture of sulfur compounds, such as sulfuric acid, sulfites, sulfates, and sulfur dioxide. Medicinally, it has assumed importance because of its widespread use in sulfa drugs and in many skin ointments. Sulfur is also employed in the production of matches, vulcanized rubber, dyes, and gunpowder. In a finely divided state and, frequently, mixed with lime, sulfur is used as a fungicide on plants. The salt, sodium thiosulfate, Na2S2O3.5H2O, commonly called hypo, is used in photography for “fixing” negatives and prints. When combined with various inert mineral fillers, sulfur forms a special cement used to anchor metal objects, such as railings and chains, in stone. Sulfuric acid is one of the most important of all industrial chemicals because it is employed not only in the manufacture of sulfur-containing molecules but also in the manufacture of numerous other materials that do not themselves contain sulfur, such as phosphoric acid.

1.4.4 Products of Sulfur

Three major product groups exist according to use: Rubber maker’s, Industrial, and Agricultural.

Rubber maker’s Sulfur

Sulfur has been used as a rubber chemical since Charles Goodyear discovered its vulcanizing properties in the mid 1800’s. Pencil erasers, rubber bumpers on automobiles, and latex gloves all use the same type of product, but in different quantities and heat variations.

Rubber maker’s sulfur products vary widely in formulation and use. Conditioning agents are added to improve flow ability, handling, and dispersion characteristics of finely ground sulfur. Oil is often added as a dust suppressant, reducing the risk of a sulfur dust explosion.

The following sulfur options are called grades, and offer a wide choice of purity, fineness, and conditioning agents for rubber processing


9

Grinding and Screening

Conditioning Agents

Oil Treatment

Industrial Sulfur

Industrial sulfurs are 99.5% minimum purity, processed into various physical shapes to provide a full range of particle sizes. This market includes pulp and paper, metals reclaiming, mining, steel, oil refining, and a multitude of other uses. Sulfur is also used in the public utilities sector as a scale inhibitor. Industrial sulfur is available as crude lumps, flakes, ground sulfur, formed pastilles, or formed briquettes. Flake sulfur can be screened to a variety of specifications.

Commercial Grades

Ground sulfurs milled to various specifications.

Arrow Roll® Refined Sulfur

Prill

Animal Feed Sulfur

Granular / Pastille

Emulsified

Flake

Agricultural Sulfur

These products are formulated for use as nutrients, soil amendments, and pesticides. Their main uses are as fungicides, insecticides, and miticides. Another common use is as a soil additive to correct alkalinity or sulfur deficiency.

Wettable Sulfurs

Wettable powders are formulated by blending dispersants and surfactants together and then milling to a very fine particle size. They can be applied as a spray or dust These products are used primarily as a fungicide or miticide. Wettable sulfur can be applied as a ground spray or aerial application. Poultry houses can be rid of depluming mites by applying spray to all interior surfaces. These products are registered by the EPA.

Flowable Sulfur


10

Dispersion or flowable type products are generally used on vine crops such as grapes, tomatoes, and peanuts. Formulated as water based dispersion weighing six pounds per gallon and is primarily used as a fungicide. This product can also be used as a soil amendment if immediate pH correction is required. These products are registered by the EPA.

Dusting Sulfur

Formulated at 98%, this product is primarily used as a fungicide. This product is registered by the EPA.

Degradable Sulfur

Formulated at 90%, this product is primarily used as a plant nutrient. Formed as a pastille or granule, degradable sulfur is also available in various sizes to conform to specific blend requirements.

1.5 Sulfur Recovery Methods

On a worldwide basis natural gas and crude oil are becoming sourer. As the sweeter, more desirable natural gas and crude oil supplies are exhausted; more and more emphasis is placed on these sour, less desirable feed stocks. The sulfur species in natural gas after its removal is generally in the form of hydrogen sulfide (H2S). The most common means of recovering the sulfur contained in hydrogen sulfide is the Clause process. This process can recover 93-99% of the sulfur contained in its feed. Recovery depends upon feed composition, age of catalyst, and number of reactor stages. The gas leaving the Clause plant is referred to as tail gas and is burnt to convert the remaining hydrogen sulfide, which is lethal at low levels, to sulfur dioxide, which has a much higher toxic limit. The off-gas stream is vented to atmosphere or sent to Tail Gas Recovery Plant.

Fig 1.3: Average production of crude oil and natural gas for sulfur extraction.

0

2000

4000

6000

8000

10000

12000

1999 2000 2001 2002 2003 2004 2005 2006 2007 2008

Pro

du

ctio

n (

MTP

D)

Sulfur Production

Production (MTPD) Crude

Production (MTPD) Natural Gas


11

Processes are differentiated on the bases of capacity as follows.

1.5.1 Medium (0.20 to 25.0 LTPD)

Sulfa Treat DO (Direct Oxidation) is a medium scale process.

1.5.2 Large (greater than 25.0 LTPD)

Clause process least expensive, well proven but only economical at large scales.

Recycle Selectox Process.

Selective Oxidation Process. - Parson’s High Activity Process. - Super Clause Process.

Wet Oxidation Based on Aqueous Solution. -Stratford and Sulfoline Process. -SulFerox Process. -Bio-SR Process.

Cold Bed Absorption Process. - CBA 4 Reactor Scheme. - CBA 3 Reactor Scheme.

Thermal Cracking of H2S.

1.5.3 Explanation of Various Processes

A brief description is given below for each of the above mentioned processes. This discussion will prove helpful in final process selection.

1.5.3.1 Sulfa Treat Direct Oxidation Process

It is a medium scale process which selectively oxidizes H2S to Sulfur and Water.

H2S + 1

2 O2 S + H2O

No equilibrium limitations are there because of good catalyst selectivity. This process recovers 90% of H2S as sulfur in a single step. It uses a patented catalyst and has a very low capital and operating costs. It can be directly operated on Natural Gas, Syngas and Hydrogen. It has got a smaller footprint than Liquid Redox or Claus Process.


12

Inlet Gas

(syngas

or natural

gas)

Liquid

knockout

Feed

Preheater

Fuel / air

Flue Gas

Air

Direct

Oxidation

Reactor

Sulfur

CondenserSulfur

To

downstream

processing

1.5.3.2 The Claus Process

This is the least expensive, well proven but only economical at large scales. It was developed by Carl Friedrich Claus in 1883. The process was later significantly modified by a German company; I. G. Farbenindustrie A. G. The Claus technology can be divided into two process steps, thermal and catalytic.

Thermal Step. In the thermal step, hydrogen sulfide-laden gas reacts in a substoichiometric combustion at temperatures above 850 °C such that elemental sulfur precipitates in the downstream process gas cooler. The H2S content and the concentration of other combustible components (hydrocarbons or ammonia) determine the location where the feed gas is burned. Claus gases (acid gas) with no further combustible contents apart from H2S are burned in lances surrounding a central muffle by the following chemical reaction:

H2S + 3

2 O2 SO2 + H2O (∆H = -4147.20 kJ/kgmol)

Gases containing ammonia, such as the gas from the refinery's sour water stripper (SWS), or hydrocarbons are converted in the burner muffle. Sufficient air is injected into the muffle for the complete combustion of all hydrocarbons and ammonia. The air to the acid gas ratio is controlled such that in total 1/3 of all hydrogen sulfide (H2S) is converted to SO2. This ensures a stoichiometric reaction for the Claus reaction (see next section below).

The separation of the combustion processes ensures an accurate dosage of the required air volume needed as a function of the feed gas composition. To reduce the process gas volume or obtain higher combustion temperatures, the air requirement can also be covered by injecting pure oxygen. Several technologies utilizing high-level and low-level oxygen enrichment are available in industry, which requires the use of a special burner in the reaction furnace for this process option.

Usually, 60 to 70% of the total amount of elemental sulfur produced in the process is obtained in the thermal process step. The main portion of the hot gas from the combustion

Fig 1.4: A typical flow

diagram of direct

oxidation (DO)

process.


13

chamber flows through the tube of the process gas cooler and is cooled down such that the sulfur formed in the reaction step condenses. The heat given off by the process gas and the condensation heat evolved are utilized to produce medium or low-pressure steam. The condensed sulfur is removed at the gas outlet section of the process gas cooler.

A small portion of the process gas can be routed through a bypass inside of the process gas cooler, as depicted in the here above mentioned figure. This hot bypass stream is added to the cold process gas through a three-way valve to adjust the inlet temperature required for the first reactor.

Catalytic Step. The Claus reaction continues in the catalytic step with activated aluminum (III) or titanium (IV) oxide, and serves to boost the sulfur yield. The hydrogen sulfide (H2S) reacts with the SO2 formed during combustion in the reaction furnace, and results in gaseous, elemental sulfur. This is called the Claus reaction:

2H2S + SO2 3

8 S8 + 2H2O (∆H = -1165.60 kJ/kgmol)

The catalytic recovery of sulfur consists of three sub steps: heating, catalytic reaction and cooling plus condensation. These three steps are normally repeated a maximum of three times. Where an incineration or tail-gas treatment unit (TGTU) is added downstream of the Claus plant, only two catalytic stages are usually installed.

The first process step in the catalytic stage is the gas heating process. It is necessary to prevent sulfur condensation in the catalyst bed, which can lead to catalyst fouling. The required bed operating temperature in the individual catalytic stages is achieved by heating the process gas in a reheater until the desired operating bed temperature is reached.

Several methods of reheating are used in industry:

Hot-gas bypass: this involves mixing the two process gas streams from the process gas cooler (cold gas) and the bypass (hot gas) from the first pass of the waste-heat boiler.

Indirect steam reheaters: the gas can also be heated with high-pressure steam in a heat exchanger.

Gas/gas exchangers: whereby the cooled gas from the process gas cooler is indirectly heated from the hot gas coming out of an upstream catalytic reactor in a gas-to-gas exchanger.

Direct-fired heaters: fired reheaters utilizing acid gas or fuel gas, which is burned substoichiometrically to avoid oxygen breakthrough which can damage Claus catalyst.

The typically recommended operating temperature of the first catalyst stage is 315 °C to 330 °C (bottom bed temperature). The high temperature in the first stage also


14

helps to hydrolyze COS and CS2, which is formed in the furnace and would not otherwise be converted in the modified Claus process.

The catalytic conversion is maximized at lower temperatures, but care must be taken to ensure that each bed is operated above the dew point of sulfur. The operating temperatures of the subsequent catalytic stages are typically 240 °C for the second stage and 200 °C for the third stage (bottom bed temperatures).

In the sulfur condenser, the process gas coming from the catalytic reactor is cooled to between 150 and 130 °C. The condensation heat is used to generate steam at the shell side of the condenser.

Before storage, liquid sulfur streams from the process gas cooler, the sulfur condensers and from the final sulfur separator are routed to the degassing unit, where the gases (primarily H2S) dissolved in the sulfur is removed.

The tail gas from the Claus process still containing combustible components and sulfur compounds (H2S, H2 and CO) is either burned in an incineration unit or further desulfurized in a downstream tail gas treatment unit.

Fig 1.5: Flow diagram for the Claus process.

1.5.3.3 Recycle Selectox Process

The Recycle Selectox Process developed by Parsons and Unocal, treats lean acid gas containing 5 to 30 mole percent H2S. The selector catalyst directly catalyzes the oxidation of H2S to SO2, eliminating the reaction furnace of Claus Process. It also catalyzes the Claus reaction of production of elemental sulfur. The exothermic Claus reaction results in a temperature increase of 30o C in first reactor stage and about 15 o C across the second stage.

The Recycle Selectox stage usually consists of one Selectox stage, followed by two Claus stages. A recycler blower dilutes the incoming acid gas with Selectox condenser.

Typical H2S conversion to sulfur is more than 80%. Total sulfur recovery with two subsequent Claus stages ranges from 94 to 96 percent. If the lean gas contains less than 5


15

percent H2S, the once-through Selectox process can be used. Except for the recycle loop, equipment arrangement is same.

Fig 1.6: Flow diagram for the Recycle Selectox Process.

1.5.3.4 Selective Oxidation Process

There are two types need to be illustrated in this account. Parson's Hi-Activity Process

In a Claus unit, complete conversion of H2S and SO2 to elemental sulfur is not possible due to limitations of thermodynamic chemical equilibrium of the Claus process. Selective oxidation of H2S to sulfur can be thermodynamically complete as indicated by the following reaction:

H2S + 1

2 O2

1

n Sn + H2O

Parson's hi-activity process utilizes a series of proprietary catalysts for direct oxidation of H2S to elemental sulfur. The Hi-activity catalysts, which are prepared with different mixtures of Iron-based metal oxides without the use of a carrier, posses low specific surfaces and wild pores , the process scheme is very similar to a conventional modified Claus unit, except that the last catalytic stage is replaced with h-activity catalyst.

Super Claus Process

The Super Claus process consists of a thermal stage followed by three of four catalytic reaction stages. The first two or three reactors are filled with standard Claus catalyst while the last reactor is filled with the selective oxidation catalyst. In the thermal stage the acid gas is burnt with a sub-stoichiometric amount of controlled combustion air such that the tail gas leaving the second reactor contains 0.80 to1.50% by volume of H2S.


16

The catalyst in the last reactor (selective oxidation reactor) oxidizes the H2S to sulfur

(H2S + 1

2 O2 S + H2O) at a very high efficiency. Because the catalyst neither oxidizes

H2S to SO2 and H2O, nor reverses the reaction of sulfur and water to H2S and SO2, a total sulfur recovery rate in the range of 99% can be obtained, depending on Claus Feed Gas composition.

1.5.3.5 Cold Bed Adsorption Process

The conventional Claus sulfur recovery process is limited by reaction equilibrium considerations to sulfur recoveries in the range of 94-97%. Very high (more than 99.8%) sulfur recoveries can be achieved by adding an amine-based tail gas cleanup process on the Claus effluent. A good example of this technology is the SCOT process licensed by Shell, which is often employed in refineries to reduce sulfur dioxide emissions to very low levels. However, amine-based tail gas cleanup units are not only expensive to build (often 80% or more of the cost of the upstream Claus plant), but expensive to operate as well. A better choice of technology for the intermediate sulfur recovery range of 98-99.5% is the so-called “sub-dew point” Claus process. This process extends the capability of the Claus process by operating the Claus reaction at a lower temperature, so that the sulfur produced by the reaction condenses. Since the Claus reaction occurs in the gas phase, this liquid sulfur does not inhibit the reaction like sulfur vapor does, resulting in a favorable shift in the reaction equilibrium and higher sulfur conversion. Amoco Corporation developed and licenses the most widely used sub-dew point Claus process.

A CBA sulfur plant consists of a conventional Claus section and a CBA section. The thermal and catalytic conversion in the conventional Claus portion of the sulfur plant usually recovers 90-95% of the inlet sulfur. Adding more conventional Claus catalytic stages beyond this point would not add much sulfur recovery because the Claus reaction is an equilibrium reaction and becomes limited by the concentrations of water and sulfur vapor in the gases flowing through the plant. The CBA portion of the sulfur plant overcomes this limitation through the use of “sub-dew point” conversion stages. Although catalytic conversion of H2S and SO2 is higher at lower reactor temperatures, conventional Claus reactors must be operated at temperatures sufficiently high to keep the sulfur produced from condensing. Sulfur catalyst will adsorb liquid sulfur in its pores, which blocks the active sites where the Claus reaction occurs. If the Claus reactor temperature is too low, the sulfur concentration in the vapor will exceed its dew point concentration, causing liquid sulfur to form and adsorb on the catalyst. Over time, this liquid sulfur will block all of the active sites in the catalyst and render the catalyst bed almost completely inactive.

A CBA reactor is operated in a cyclic fashion to avoid complete catalyst deactivation from liquid sulfur blocking the active sites. The CBA reactor is operated at low temperature (250-300°F/120-150°C) initially so that it is below the sulfur dew point of the reaction products (i.e., “sub-dew point”) and the sulfur formed is condensed and adsorbed on the catalyst. After operating in this manner for a period of time, the CBA reactor is “regenerated” by flowing hot gas through the reactor to vaporize the adsorbed liquid sulfur,


17

which is then condensed and removed in a down-stream sulfur condenser. This process is analogous to the processing steps used when dehydrating gas streams with molecular sieves. There are normally two or more CBA reactors in series so that at least one can be operating sub-dew point while the other is being regenerated. Not only does a CBA reactor benefit from a more favorable Claus reaction constant at its lower operating temperature, it also has the advantage of shifting the Claus reaction equilibrium. The Claus reaction is a vapor-phase reaction, so condensing the sulfur product removes it from the vapor, forcing the equilibrium in the Claus reaction further to the right, toward higher conversion. These two factors allow much higher sulfur conversion than in a conventional Claus reactor, resulting in overall sulfur recovery efficiencies in excess of 98-99.5% for CBA plants.

The cyclic nature of the CBA process requires process gas switching valves that must perform in very demanding sulfur vapor services. This has caused significant operation and maintenance problems in CBA plants designed by other engineering companies and contractors.

1.5.3.6 Thermal Cracking of H2S

In this process operation at significantly high temperatures is made possible and economical by oxidation of part of the H2S to provide the energy required for the decomposition reaction to proceed to a significant extent. Partial oxidation of H2S in the H2S-containing fuel gas is carried out in the presence of an inert, porous, high-capacity medium and the intense heat exchange results in flame temperatures that significantly exceed the adiabatic flame temperature of the gas mixture. By coupling the partial oxidation of H2S in the porous medium with the H2S decomposition, temperatures as high as 1400°C (1673K) can be achieved economically within a reaction zone without the input of external energy, and therefore, no additional CO2 emissions. In this reaction zone, the self-sustaining conditions are very favorable for the decomposition reaction to proceed to an industrially significant extent, within a slowly propagating thermal wave.

Fig 1.7: Thermal cracking of H2S.


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1.6 The Claus Process

The Claus process is the most significant gas desulfurizing process, recovering elemental sulfur from gaseous hydrogen sulfide. First invented over 100 years ago, the Claus process has become the industry standard.

1.6.1 History

The process was invented by Carl Friedrich Claus, a chemist working in England. A British patent was issued to him in 1883. The process was later significantly modified by a German company called I. G. Farbenindustrie A. G.

1.6.2 Description

The multi-step Claus process recovers sulfur from the gaseous hydrogen sulfide found in raw natural gas and from the by-product gases containing hydrogen sulfide derived from refining crude oil and other industrial processes. The by-product gases mainly originate from physical and chemical gas treatment units (Selexol, Rectisol, Purisol and amine scrubbers) in refineries, natural gas processing plants and gasification or synthesis gas plants. These by-product gases may also contain hydrogen cyanide, hydrocarbons, sulfur dioxide or ammonia.

Gases with an H2S content of over 25% are suitable for the recovery of sulfur in straight-through Claus plants while alternate configurations such as a split-flow set up or feed and air preheating can be used to process leaner feeds.

Hydrogen sulfide produced, for example, in the hydro desulfurization of refinery naphthas and other petroleum oils, is converted to sulfur in Claus plants The overall main reaction equation is:

2H2S + O2 S2 + 2H2O

In fact, the vast majority of the 64,000,000 metric tons of sulfur produced worldwide in 2005 was byproduct sulfur from refineries and other hydrocarbon processing plants. Sulfur is used for manufacturing sulfuric acid, medicine, cosmetics, fertilizers and rubber products.

Inevitably a small amount of H2S remains in the tail gas. This residual quantity, together with other trace sulfur compounds, is usually dealt with in a tail gas unit. The latter can give overall sulfur recoveries of about 99.8%, which is very impressive indeed.

Gases containing ammonia, such as the gas from the refinery's sour water stripper (SWS), or hydrocarbons are converted in the burner muffle. Sufficient air is injected into the muffle for the complete combustion of all hydrocarbons and ammonia. Air to the acid

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gas is controlled such that in total 1/3 of all hydrogen sulfide (H2S) is converted to SO2. This ensures a stoichiometric reaction for the Claus reaction.

Fig 1.8: The Claus process for sulfur recovery.

1.6.3 Simplified Process Description

The hot combustion products from the furnace at 1000- 1300 °C enter the waste heat boiler and are partially cooled by generating steam. Any steam level from 3 to 45 bar g can be generated.

The combustion products are further cooled in the first sulfur condenser, usually by generating LP steam at 3 – 5 bar g. This cools the gas enough to condense the sulfur formed in the furnace, which is then separated from the gas and drained to a collection pit.

In order to avoid sulfur condensing in the downstream catalyst bed, the gas leaving the sulfur condenser must be heated before entering the reactor.

The heated stream enters the first reactor, containing a bed of sulfur conversion catalyst. About 70% of the remaining H2S and SO2 in the gas will react to form sulfur, which leaves the reactor with the gas as sulfur vapor.

The hot gas leaving the first reactor is cooled in the second sulfur condenser, where LP steam is again produced and the sulfur formed in the reactor is condensed.

A further one or two more heating, reaction, and condensing stages follow to react most of the remaining H2S and SO2.

The sulfur plant tail gas is routed either to a Tail Gas treatment Unit for further processing, or to a Thermal Oxidizer to incinerate all of the sulfur compounds in the tail gas to SO2 before dispersing the effluent to the atmosphere.

1.6.4 Process Improvements


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Over the years many improvements have been made to the Claus process. Recent developments include:

SUPERCLAUS (TM). A special catalyst in the last reactor oxidizes the H2S selectively to sulfur, avoiding formation of SO2. Significantly higher conversions are obtained at modest cost.

Oxygen Claus. The combustion air is mixed with pure oxygen. This reduces the amount of nitrogen passing through the unit, making it possible to increase throughput.

Better Catalysts. Higher activities have been achieved with catalysts that provide higher surface areas and macro porosity.

More improvements can be expected. Here are some possibilities.

CS2 destruction. Carbon disulfide (CS2) is a side product made in the furnace. Laboratory work has shown that special catalysts operating in the furnace can destroy the CS2 before it gets into the catalytic section. A commercially available catalyst like this might be developed for use in a Claus plant.

Catalyst Temperature Policy. The conversion of H2S goes faster at higher temperatures, but a more favorable equilibrium is obtained at lower temperatures. It isn't obvious whether higher or lower temperatures are needed in the third converter. Kinetic modeling may supply the answer, thereby improving conversion or reducing catalyst replacement cost.

1.6.5 Claus Process Auxiliaries

Following are some of the auxiliaries being used in the Claus process.

1.6.5.1 Blow Down System

Boiler blow down flows is collected and drained into SBD-1. Steam Blow Down Drum. The steam is vented from the top of SBD-1 and liquid flows from the bottom to SBC-1,Steam Blow Down Cooler. The blow down flows from E-2 to the drain system.

1.6.5.2 Fuel System

Fuel gas is supplied from OSBL services. The three users are F-1 (Muffle furnace), TG-1 ( Tail Gas Incinerator), and a PA-1(Package Auxiliary Box). In normal operation, fuel flows only to TG-1. The fuel to F-1 is used only in start up when heating the unit. The fuel to PA-1 is used only when the clause process is shut down, then PA-1 is used to supply steam for heating services, primarily on the liquid sulfur containing lines and equipment.

There are two different fuel gases. One is natural gas types and is supplied to all three-fuel users. The other fuel gas is vaporized LPG which is supplied only to PA-1 to provide an alternate fuel for this equipment.

Chapter 2 MAJOR EQUIPMENTS USED IN CLAUS PROCESS

2.1 Introduction

he major equipment items used in the project are discussed in this chapter, generally in the order of process flow through the SRU. The concept presented is intended to improve the SRU reliability and are not intended to be complete design guidelines. A general comment that applies to all equipment items is to provide pressure point/sample point connections between all major equipment

items. These will prove invaluable in troubleshooting the SRU.

2.2 Reaction Furnace (F-100)

The main burner and reaction furnace combine to form the SRU thermal reactor. The burner and reaction furnace are normally mounted horizontally with the burner coaxially mounted on the end of the reaction furnace.

The thermal reactor is the heart of the SRU even though it is frequently selected or designed without considering its level of importance. We consider the main burner to be the most important piece of equipment in the SRU. The burner must perform the function of burning one third of the feed hydrogen disulfide (H2S) to sulfur dioxide (SO2) to satisfy the stoichiometric requirements of the modified Claus process, while also destroying impurities in the acid gas feed and consuming all of the oxygen in the combustion air. The burner must be capable of performing efficiently at normal operating feed rates and low turn down rates. The burner must also be capable of substoichiometric burning of natural gas during start up and shut down operations.

The use of a very efficient mixing, high intensity burner is preferred. Inefficient

burners are frequently employed in SRU’s around in many industries. Frequently, these

burners are not able to achieve adequate destruction of impurities, complete oxygen

consumption and tend to produce some amounts of sulfur trioxide (SO3). These

shortcomings can result in equipment corrosion, catalyst deactivation, and plugging of

piping, equipment and catalyst beds. All of these adverse results reduce the SRU reliability

since under these conditions the requirement for shutting down of unit increases for

maintenance repairs, catalyst change and/or unblocking an obstruction. The burner must

accomplish the combustion reactions. The combustion reactions are relatively fast. The

reaction furnace provides the residence time at high temperature required

T

Chapter 2 Major Equipments Used in the Claus Process

22

for the Claus reactions and side reactions to occur. Many feed impurities and intermediate

components must be destroyed in the reaction furnace or they will cause downstream

problems.

These components must have adequate time for the reactions to reach

completion/equilibrium. The reaction furnace should have 5 to 7 seconds residence time.

The specific features and residence time required for an individual reaction furnace are

dependent on several factors including the operating temperature and expected feed

impurities.

2.3 Waste Heat Boiler (B-100)

A WHR boiler is a closed vessel in which water or other fluid is heated. The heated or

vaporized fluid exits the WHR boiler for use in various processes or heating applications.

The various types of waste heat boilers include:

• Fire-Tube Boiler.

• Water-Tube Boiler.

• Vertical boiler.

• Hydronic Boiler.

Fire-Tube Boiler:

A fire-tube boiler is a type of boiler in which hot gases from a fire pass through one or more tubes running through a sealed container of water. The heat energy from the gases passes through the sides of the tubes by thermal conduction, heating the water and ultimately creating steam.

Water-Tube Boiler:

It is a type of boiler in which water circulates in tubes heated externally by the fire.

Water tube boilers are used for high-pressure boilers. Fuel is burned inside the furnace,

creating hot gas which heats water in the steam-generating tubes. In smaller boilers,

additional generating tubes are separate in the furnace, while larger utility boilers rely on

the water-filled tubes that make up the walls of the furnace to generate steam.

Vertical Boiler:

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The Cyclone Hot Water Boilers provide for exceptionally high efficiencies, lower fuel

costs, and extremely rugged construction. Compact space saving vertical design four-pass

design shock proof, no tubes to loosen or burn out. Convenient access to "eye high" burner

solid state controls for trouble free operation factory assembled, fully automatic UL and

ASME CSD-1. Simple and inexpensive to install.

Hydronic Boiler:

Hydronic boilers are used in generating heat for residential and industrial purposes. They are the typical power plant for central heating systems fitted to houses in northern Europe (where they are commonly combined with domestic water heating), as opposed to the forced-air furnaces or wood burning stoves more common in North America. The Hydronic boiler operates by way of heating water/fluid to a preset temperature (or sometimes in the case of single pipe systems, until it boils and turns to steam) and circulating that fluid throughout the home typically by way of radiators, baseboard heaters or through the floors. The fluid can be heated by any means...gas, wood, fuel oil, etc, but in built-up areas where piped gas is available, natural gas is currently the most economical and therefore the usual choice. The fluid is in an enclosed system and circulated throughout by means of a motorized pump.

2.4 Sulfur Condensers (E-100, E-102, E-104, E-106)

The Claus sulfur recovery process consists of four repeating steps for sulfur

condensation. Sulfur condensers serve the primary function of cooling and condensing

sulfur formed in the upstream reaction step. Sulfur condensers are normally horizontal,

kettle type shell and tube boilers. However, sulfur condensers are unique heat exchangers.

In addition to condensing product sulfur from the process gases, the liquid sulfur must also

be separated from the process gases before they flow to the next processing step. This is

normally done in an oversized outlet channel. Sulfur condensers are also unique because

the process gas flow rate through the condensers must be maintained within a specific

operating range/velocity or there will be adverse effects on the process. The term to

describe this flow property is the “Mass Velocity”, which is normally expressed as “pounds

of process gas flow per second per square foot of cross sectional flow area”. The

recommended mass velocity operating range is 1.50 to 5.50 lb/sec-ft2.

Ideally sulfur is condensed from the process gas at the cool condenser tube walls,

flows from the tube into the outlet channel, is separated from the process gas, and is

drained from the condenser. If the mass velocity is too high, liquid sulfur can be entrained in

the process gas and be carried to the next stage or from the SRU instead of draining from

the tube. If the mass velocity is too low, sulfur can condense in the vapor as a very small

droplet or fog. The sulfur fog droplets are so small that the gas stream carries them much

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like atmospheric fog or smoke to the next stage or from the SRU. In either of these cases,

sulfur recovery is lost. Lower recovery does not directly affect SRU reliability, but lower SRU

recovery will cause additional load on the downstream tail gas cleanup unit or increase the

plant emissions. Either of these conditions may ultimately cause the SRU to be shut down

permanently for maintenance.

Some designs utilize one or more of condensers as BFW preheaters upstream of the

WHB. The lower level heat from the condenser is used to increase the generation of higher

pressure stream by preheating the WHB feed water. The condenser shell must operate at

the BFW header pressure with this design. But this will exert excessive stress on the tube to

tube sheet attachment and reduces unit reliability. If sulfur condensers are used as BFW

preheaters, the tubes should be strength welded to the tube sheet. Some designs allow the

first condenser to act like the BFW preheater but some steam is generated on the shell side

of the sulfur condenser. Because the shell is not designed for vaporizing conditions, vapor

blanketing of some tubes can occur. This can result in overheating the tubes and sulfide

corrosion.

Some designs also use cold BFW on the shell of the final sulfur condenser to

minimize the process outlet temperature and maximize sulfur recovery.

As mentioned above, generation of low pressure steam to minimize the process

outlet temperature is preferred because if the BFW is too cold, there is a potential to freeze

the sulfur in the tubes.

It is essential for each sulfur condenser to have an independent sulfur seal and look

box. The ability to observe the sulfur production from each condenser is a very valuable

process evaluation and troubleshooting tool. The sulfur rate, consistency of rate, color,

temperature, and presence of bubbles are all important information items that can only be

obtained from individual seals and look boxes.

Each sulfur condenser drain line, sulfur seal, look box and rundown line to the sulfur

pit should be fully steam jacketed. The drain line between the condenser and seal should

have a steam jacketed plug valve located as close as practical to the condenser to allow on-

line rodding of the drain line and sulfur seal. Clear access must be provided for rodding the

drain line and overhead access must be provided to rod the seal.

Some plants have implemented a method to flush sulfur seals to keep the seals open

and free flowing. Steam jacketed piping with block valves are employed from the sulfur

pump discharge to the inlet of each sulfur seal. This allows flushing the individual seals by

closing the block valve in the drain line from the condenser and flowing product sulfur from

the pump discharge through the seal and back to the sulfur pit. This is an excellent and safe


25

method to keep the seals free flowing. Some plants use steam to periodically blow the

sulfur seals when there is an indication of partial plugging. While this method normally

works, we feel it should not be done as a routine practice because of the safety risks from

the hot liquid sulfur.

2.5 Heaters (E-101, E-103, E-105)

Reheaters definitely offer more options to the process designer than any other item

in the SRU. There are two general types of reheaters, direct and indirect. There are also

multiple options within each type.

Indirect method is preferred over direct method; however, each method has specific

applications where it should be considered. Each reheating method is briefly discussed

below.

2.5.1 Direct Reheat Methods

Direct reheat methods use a hot gas stream that is mixed with the process gas to increase

the temperature of the mixed stream to the desired inlet temperature of the downstream catalytic

reactor. The hot gas stream may originate within the process or from combustion. The direct reheat

methods are hot gas bypass, acid gas fired line burner, and natural gas fired line burner, and natural

gas fired line burner, if any of the direct methods are used, it is very important to insure there is

adequate mixing of the streams upstream of the temperature control point.

2.5.2 Hot Gas Bypass

This method has been used in many SRU's. It uses a hot stream from the first pass outlet of

the WHB (1000-1200F) to mix with process gas streams from the sulfur condensers.

It is inexpensive to install, but it has the disadvantages of lowering sulfur recovery by

bypassing conversion steps with a portion of the process gas, poor turndown performance, and high

temperature sulfide corrosion of carbon steel piping and control valves. The corrosion problems can

be minimized with proper metallurgy, but this is often not done because the cost is higher, and low

cost is a primary reason to use hot gas bypass reheat.

The only reason to use hot gas bypass reheat in current design is for very small, isolated

location plants that do not have access to high pressure steam or adequate, reliable electric power

supplies.

2.5.3 Acid Gas Fired Line Burner


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Acid gas fired burners have been used in many SRU’s. There primary advantage the ability to

achieve any desired catalytic reactor inlet temperature. However, line burners have disadvantages.

The overall sulfur recovery is normally reduced because acid gas bypasses some conversion steps.

The burner air/fuel ratio must be closely controlled or oxygen breakthrough, soot formation, and/or

SO3 formation is likely.

We prefer to use a steam reheater design in which the high pressure steam is on the tube

side of a U-tube type heat exchanger. This type design avoids having to design the shell for the high

pressure steam and avoids tube to tube sheet stresses which can cause failures with steam leakage

into the process and SRU shutdown to make repairs. The U-tube bundle, free to expand within the

shell, avoids these mechanical stresses.

2.6 Catalytic Reactors (R-100, R-101, R-102)

Reactor is a vessel in which different species react to forma product under specified operating conditions.

TYPES OF RECTOR:

Chemical reactors come in the form of vessels or tanks for batch reactors or back-mix flow reactors ,as cylinders for fluidized bed reactors or as single or multiple tubes inside a cylindrical container for plug flow reactor.

CATALYTIC REACTOR:

Use of catalysts requires modification to basic reactor design in order to account for mass and energy transport issues arising from catalysts.

FIXED BED REACTOR (FBR):

These reactors are solid-catalyst containing vessels. Their design can lead to high pressure drops. These units are generally used in heterogeneous catalysis where the catalysts and reacting species are of different phases. The major advantage of such units is their simplicity and ease of catalyst access for maintenance and regeneration. Use of multiple fixed beds can improve both heat transport and control resulting in improved performance while maintaining the relative simplicity of this reactor arrangement.

MULTIPLE TUBULAR REACTORS:

These types of reactors are modified multiple fixed bed units , where the multiple beds are catalyst-filled tubes arranged in parallel with a heat conducting fluid flowing outside the tubes. These reactors offer good thermal control and uniform residence time distribution , but experience increased complexity as well as catalyst in accessibility. Catalyst access is somewhat simplified by packed tube arrangement although packing and removing the catalyst from the tubes can still be difficult.

SLURRY REACTOR:


27

Reaction of slurries containing solid particles that can be physically separated from the suspension fluid are often best performed in agitated tank-type fluid reactors. The reactor offer simplicity good transport properties and control while sacrificing nothing in catalyst access since catalyst particles can be added and removed continuously. There is however, an increased element of equipment degradation due to particle impingement on the fluid handling equipment, such as impellers, nozzles and pipes.

MOVING BED REACTOR:

These units are also fluid reactors used where the fluid contains solid particles that can be physically separated from from the suspension fluid. In this case however, the slurry travels through the reactor in essentially plug flow. Again simplicity , access and control are good with a uniform residence time distribution.

FLUIDIZED BED REACTORS:

These are reactors with a gas phase-working fluid that requires gas flow around and past fine particles at a rate sufficient to fluidize the particles suspended within the reactor. There are considerable operating difficulties associated with initiating and running fluidized bed reactors due to flow and suspension issues. Further these types of reactors have large residence time distribution of the ease of back flow in the gas and approach CSTR behaviour. The advantages of these reactors are their ability to process fine particles and suitability to high reaction rate processes.

THIN OR SHALLOW BED REACTORS:

These designations are reserved for reactors where the reactant fluid flow through catalyst meshes or thin beds. These are simple reactors particularly suitable for fast reaction that require good control where catalyst access is important for purposes of catalyst reactivation or maintenance or where large heats of reaction are involved.

DISPERSION REACTORS:

These types of reactors are fluid-containing vessels that allow dispersion of liquid and gas phase reactants by bubbling the latter through the liquid or dripping the liquid into the gas stream or into a less dense liquid, to achieve increased contact area and reaction performance. Even though these reactors are simple and inexpensive reactors, they require careful planning due to their sensitivity to flow behaviour.

FILM REACTORS:

A reactor design that maximizes contact area for gas/liquid reactions is film the reactor that brings together a gas and liquid as a thin film over a solid support. This type of reactor offers an added benefit of increased thermal control via the solid support. Such as arrangement also allows for complex phase dependent reactions in which solid, liquid and gas phase are involved.


28

SELECTION OF REACTORS:

The selection of best reactor type for a given process is subjected to # of major consideration. Such design aspects, for example,

1) Temperature and pressure of the reaction. 2) Need for removal or addition of reactants and products. 3) Required pattern of product delivery (continuous or batch wise) 4) Catalyst use consideration, such as the requirements for solid catalyst particle

replacement and contact with fluid reactants and products; 5) Relative cost of reactors.

REACTOR USED FOR CATALYTIC STEP OF CLAUSE PROCESS:

The reactor used for catalytic step of clause process of sulphur recovery is the fixed bed catalytic reactor. The most important characteristic of FBR is that material flows through the reactor as plug, all of the stream flows at the same velocity, parallel to reactor axis with no back mixing. All material present at any given reactor cross-section has had an identical residence time.

They can be classified according to the manner in which the temperature is controlled into reactors with adiabatic reaction control.

Fixed bed reactors contain a bed of catalyst pellets. The catalyst lifetime in these reactors is greater than three months. These are rectors widely used in petro-chemical industries. They can generally be carried out continuously at low to medium pressure.

Fixed bed reactors are often referred to as packed bed reactors. They may be regarded as the workhorse of the chemical industry w.r.t. number of reactors employed and the economic value of materials produced. In a FBR, for a fluid- solid reaction, the solid catalyst is present as a bed of relatively small particles randomly oriented and fixed in position. The fluid moves by convective flow through the spaces between the particles. There may also be diffusive flow or transport within the particles.We also focus on steady-state operation thus ignoring any implications of catalyst deactivation with time.

INDUSTRIAL IMPLICATIONS OF FBR:

Synthesis of Ammonia Production of styrene monomer by dehydrogenation of ethyl benzene. Alkylation of benzene to ethyl benzene. Production of sulphuric acid. Synthesis of butynediol from acetylene and formaldehyde.

FLOW ARRANGEMENT:

Traditionally most FBR are operated with axial flow of liquid down the bed of solid.


29

ADIABATIC MODE OF OPERATION:

In adiabatic operation, no attempt is made to adjust temp within the bed by means of beat transfer. For a reactor consisting of one bed of catalyst, this defines the situation thermally. If catalyst is divided into two or more beds arranged in series (a multistage reactor). There is an opportunity to adjust temp b/w stages, even if each step is operated adiabatically this may be done in two ways:

(1) Ist involves the inter-stage beat transfer by means of heat exchangers used for either exothermic or endothermic reaction.

(2) Second called the COLD-SHOT COOLING, can be used for exothermic reactions

PURPOSE OF ADJUSTMENT OF TEMP: (1) To shift an equilibrium limit so as to increase fractional conversion or yield. (2) To maintain relatively light rate of reaction to decrease amount of catalyst and

size of vessel required.

DESIGN CONSIDERATION:

The most important factor to be considered in the design of such reactors is: Residence time distribution: influence on conversion and selectivity. Temp control: maintenance of temp limits, axially and radialy , min temp diff,

b/w reactor medium and catalyst surface , as well as within the catalyst particle. Catalyst lifetime and catalyst regeneration Pressure drop as a function of catalyst shape and gas velocity. In addition to flow, thermal and bed arrangement an important design consideration

is the amount of catalyst required and its possible distribution over two or more stages. This is the measure of size of reactor. The depth (L) and diameter (D) of each stage must also be determined.

CATALYST USED:

Catalyst used in catalyst step of clause process is activated alumina (Al203) in the form of spherical pellets.

CONSIDERATION OF PARTICLE AND BED CHARACTERISTICS:-

Characteristics of a catalyst particle include its chemical composition, which primarily determines its catalyst activity and its physical properties such as size, shape, density and porosity or voidage which determines the diffusion characteristics.


30

2.7 Sulfur Pits

Product sulfur is normally collected in a below grade, concrete pit equipped with steam coils

to keep the sulfur molten. The pit doesn't directly effect the SRU process operation until the SRU

must be shut down because of problems with the pit. Some common sulfur pit problems are steam

coil leakage, sulfur pump failure, internal sulfur fires, and even internal explosions. There are a few

design features that will significantly improve the reliable operation of the sulfur pit.

1. Construct the pit using sulfate resistant concrete with limestone- free aggregate.

2. Use alloy piping for the steam coil steam supply down comers and condensate risers,

and any internal components such as ladder rungs that will be alternately covered with

liquid sulfur and then exposed to air as the pit level changes.

3. Install dual steam jacketed sulfur transfer pumps.

4. Use a fully steam jacketed steam eductor to continuously draw atmospheric air into the

pit, sweeping vapor space to prevent the accumulation of H2S.

5. Steam snuffing connection(s) for extinguishing internal sulfur fires.

6. The number of inlets depends on the size and configuration of the pit.

Chapter 3 MATERIAL BALANCE OF THE SULFUR RECOVERY UNIT

(SRU)

3.1 Introduction

he material balance across the proposed sulfur recovery unit (SRU) is done by the conservation equation of mass, as is done conventionally. A system must be defined to account for the streams entering and leaving. In our case the obvious selection is the sulfur recovery unit (SRU) itself, while all the other premises are considered surroundings. Some preliminary bases are to be specified for the sake of convenience in the calculations. Following

specifications are taken to meet the above mentioned situation:

Sulfur production: 80 tons per day

Time of operation: 1 hr

Now the material balance calculations are made first along the whole unit and then across individual equipments. It is to be noted that whether the calculations are made across the whole unit or the individual equipments, the basic law of conservation of mass equation remains the same and is given as:

Amount of substance Amount of substance Amount of substance

entering the system - leaving the system + generated within the -

through the boundaries through the boundaries system boundaries

Amount of substance Amount of substance

consumed within the = accumulated within the (3.1)

system boundaries system boundaries

3.2 Overall Material Balance

The chemical reactions taking place are:

T

Chapter 3 Material Balance of the SRU

32

Main Reactions:

1- H2S + 3

2 O2 SO2 + H2O (3.2)

2- SO2 + 2H2S 3

8 S8 + 2H2O (3.3)

Side Reactions:

3- CH4 + 2O2 CO2 + 2H2O (3.4)

4- C2H6 + 𝟕

2 O2 2CO2 + 2H2O (3.5)

5- C3H8 + 5O2 3CO2 + 4H2O (3.6)

Fig 3.1: Overall material balance across SRU.

NOTE: In all the diagrams the compositions are mentioned in mole fraction basis

Sulfur production target (S8) = 80 ton/day

=

= 13.0 kgmol/hr

H2S required by S8 = 16

3× 13.0 kgmol (from equ. 3.3)

80 tons

1 day

1000 kg

1 kgmol

day 24 hr

1 ton

256.5 kg

Flow Rate = ?

H2S = 0.523

CO2 = 0.376

CH4 = 0.004

H2O = 0.010

C2H6 = 0.0008

C3H8 = 0.0001 Flow Rate = ?

O2 = 0.210

N2 = 0.790

Flow Rate = 13 kgmol/hr

S8 = 0.999

H2O = 0.001


33

= 69.30 kgmol

H2S supply for S8 (on account for 99.9% conversion from the Claus process)

= 0.999 × 69.30 kgmol = 69.37 kgmol

SO2 required by S8 = 8


= 34.65 kgmol

H2S consumed for SO2 = 34.65 kgmol (from equ. 3.2)

Total H2S supplied = 69.37 kgmol + 34.65 kgmol = 104.03 kgmol

Total acid gas feed supply = 104.03 kgmol /0.523 = 198.80 kgmol

Now the Stream-01 composition is as follows:

Stream-01 composition

Component Mole fraction Flow rate-F01 (kgmol/hr)

H2S 0.523 104.03

CO2 0.376 74.83

CH4 0.004 0.95

H2O 0.010 18.80

C2H6 0.0008 0.16

C3H8 0.0001 0.02

TOTAL 1.0 198.80

3.3 Material Balance across Furnace F-100

It is to be noted that according to the specifications of the Claus process only 30% of total sulfur dioxide produces in the furnace is converted into elemental sulfur.1

Fig 3.2: Material balance across furnace F-100. 1 “Sulfur Recovery”, GPSA Engineering data book Vol. 2, 11

th edition, 1998. Chapter 22

Flow Rate = 198.80 kgmol/hr

H2S = 0.523

CO2 = 0.376

CH4 = 0.004

H2O = 0.010

C2H6 = 0.0008

C3H8 = 0.0001


O2 = 0.210

N2 = 0.790

Flow Rate = ?

H2S = ?

CO2 = ?

N2 = ?

SO2 = ?

H2O = ?

S8 = ?


34

In the furnace the following chemical reactions are taking place:

Main Reactions:

1- H2S + 3

2 O2 SO2 + H2O (3.2)

2- SO2 + 2H2S 3

8 S8 + 2H2O (3.3)

Side Reactions:

3- CH4 + 2O2 CO2 + 2H2O (3.4)

4- C2H6 + 7

2 O2 2CO2 + 2H2O (3.5)

5- C3H8 + 5O2 3CO2 + 4H2O (3.6)

SO2 produced = 34.65 kgmol (from equ. 3.3)

H2S still available for S8 production = 104.03 kgmol - 34.65 kgmol = 69.37 kgmol

H2S consumed for S8 production = 10.4 kgmol × 2 = 20.80 kgmol

H2S remaining = 69.37 kgmol - 20.80 kgmol = 48.58 kgmol

SO2 consumed = 34.65 kgmol × 0.30 = 10.40 kgmol

SO2 remaining = 34.65 kgmol - 10.40 kgmol = 24.25 kgmol

S8 produced = 3

8×10.40 kgmol = 3.90 kgmol

O2 required in SO2 formation = 52.0 kgmol (from equ. 3.2)

O2 required in CH4 combustion = 1.90 kgmol (from equ. 3.4)

O2 required in C2H6 combustion = 0.55 kgmol (from equ. 3.5)

O2 required in C3H8 combustion = 0.10 kgmol (from equ. 3.6)

Total O2 required = 52.0 kgmol + 1.90 kgmol + 0.55 kgmol +

0.10 kgmol

= 54.54 kgmol

Air fed to furnace = 54.54 kgmol / 0.210

= 259.70 kgmol

N2 going in = N2 going out = 259.70 kgmol × 0.790


35

= 205.20 kgmol

CO2 generated in CH4 combustion = 0.95 kgmol (from equ. 3.4)

CO2 generated in C2H6 combustion = 0.31 kgmol (from equ. 3.5)

CO2 generated in C3H8 combustion = 0.05 kgmol (from equ. 3.6)

CO2 going out = 0.95 kgmol + 0.31 kgmol + 0.05 kgmol +

74.83 kgmol

= 76.16 kgmol

H2O formed in SO2 production = 34.65 kgmol (from equ. 3.2)

H2O formed in CH4 combustion = 1.90 kgmol (from equ. 3.4)

H2O formed in C2H6 combustion = 0.47 kgmol (from equ. 3.5)

H2O formed in C3H8 combustion = 0.08 kgmol (from equ. 3.6)

H2O formed in S8 production = 20.80 kgmol (from equ. 3.3)

Total H2O produced = 34.65 kgmol + 1.90 kgmol + 0.47 kgmol +

0.08 kgmol + 20.80 kgmol

= 58.0 kgmol

H2O going out = 58.0 kgmol + 18.80 kgmol = 76.72 kgmol

Now the Stream-02, Stream-03 and Stream-04 compositions are as follows:



O2 0.210 54.54

N2 0.790 205.20

TOTAL 1.0 259.70



H2S 0.116 48.58

CO2 0.183 76.16

N2 0.50 205.20

SO2 0.058 24.25


36

H2O 0.140 58.0

S8 0.010 3.90

TOTAL 1.0 416.0

There is no need for the calculation of material balance across the waste heat boiler B-100 since no material change takes place in there. So the composition of Stream-03 and Stream-04 are identical.



H2S 0.116 48.58

CO2 0.183 76.16

N2 0.50 205.20

SO2 0.058 24.25

H2O 0.140 58.0

S8 0.010 3.90

TOTAL 1.0 416.0

Only energy changes occur and in the subsequent chapter related to the energy balance, calculations are made across it.


All of the sulfur produced in the furnace F-100 is condensed in the first condenser E-100, along with some water. The purity of sulfur extracted is 99.9%.

Fig 3.3: Material balance across condenser E-100.

S8 going in Stream-S21 = 3.90 kgmol


H2S = 0.116

CO2 = 0.183

N2 = 0.50

SO2 = 0.058

H2O = 0.140

S8 = 0.010

Flow Rate = ?

H2S = ?

CO2 = ?

N2 = ?

SO2 = ?

H2O = ?

Flow Rate = ?

S8 = 0.999

H2O = 0.111


37

Total amount of Stream-S21 = 3.90 kgmol

0.999 = 3.91 kgmol

H2O going in Stream-S21 = 3.91 kgmol × 0.111 = 0.004 kgmol

H2O going in Stream-05 = 58.0 kgmol – 0.004 kgmol = 57.90 kgmol

Now the Stream-S21 and Stream-05 compositions are as follows:

Stream-S21 composition

Component Mole fraction Flow rate-FS21 (kgmol/hr)

S8 0.999 3.90

H2O 0.001 0.004

TOTAL 1.0 3.904



H2S 0.117 48.58

CO2 0.184 76.16

N2 0.50 205.20

SO2 0.058 24.25

H2O 0.140 57.90

TOTAL 1.0 412.0

Again there is no need for the application of material balance calculations around the heat exchanger E-101. So the Stream-06 has the same composition as that of Stream-05.



H2S 0.117 48.58

CO2 0.184 76.16

N2 0.50 205.20

SO2 0.058 24.25

H2O 0.140 57.90

TOTAL 1.0 412.0



38

Fig 3.4: Material balance across reactor R-100.

Now, according to the specifications of the Claus process, the reactor R-100 converts only 70% of the incoming sulfur dioxide into elemental sulfur.1 Thus:

SO2 consumed in S8 production = 24.25 kgmol × 0.70 = 16.90 kgmol

SO2 remaining = 24.25 kgmol – 16.90 kgmol = 7.27 kgmol

H2S consumed in S8 production = 16.90 kgmol × 2 (from equ. 3.3)

= 33.90 kgmol

H2S remaining = 48.58 kgmol – 33.90 kgmol = 14.62 kgmol

H2O formed along with S8 = 16.90 kgmol × 2 (from equ. 3.3)

= 33.90 kgmol

H2O going out of reactor = 57.90 kgmol + 33.90 kgmol = 91.86 kgmol

S8 produced = 3


= 6.36 kgmol




H2S 0.036 14.62

CO2 0.190 76.16

N2 0.511 205.20

SO2 0.018 7.27

1 “Sulfur Recovery”, GPSA Engineering data book Vol. 2, 11



H2S = 0.117

CO2 = 0.184

N2 = 0.50

SO2 = 0.058

H2O = 0.140

Flow Rate = ?

H2S = ?

CO2 = ?

N2 = ?

SO2 = ?

H2O = ?

S8 = ?


39

H2O 0.228 91.86

S8 0.015 6.36

TOTAL 1.0 401.50


As in the previous case all of the produced sulfur is condensed through the condenser and then withdrawn from the collecting pits.

Fig 3.5: Material balance across condenser E-102



0.999 = 6.37 kgmol


H2O going in Stream-08 = 91.86 kgmol – 0.006 kgmol = 91.85 kgmol




S8 0.999 6.37

H2O 0.001 0.0067

TOTAL 1.0 6.38


H2S = 0.036

CO2 = 0.190

N2 = 0.511

SO2 = 0.018

H2O = 0.228

S8 = 0.015

Flow Rate = ?

H2S = ?

CO2 = ?

N2 = ?

SO2 = ?

H2O = ?

Flow Rate = ?

S8 = 0.999

H2O = 0.111


40



H2S 0.037 14.62

CO2 0.192 76.16

N2 0.520 205.20

SO2 0.018 7.27

H2O 0.232 91.86

TOTAL 1.0 395.12




H2S 0.037 14.62

CO2 0.192 76.16

N2 0.520 205.20

SO2 0.018 7.27

H2O 0.232 91.86

TOTAL 1.0 395.12


Fig 3.6: Material balance across reactor R-101


H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232

Flow Rate = ?

H2S = ?

CO2 = ?

N2 = ?

SO2 = ?

H2O = ?

S8 = ?


41





= 11.64 kgmol



= 11.64 kgmol


S8 produced = 3


= 2.18 kgmol




H2S 0.007 2.98

CO2 0.194 76.16

N2 0.524 205.20

SO2 0.003 1.45

H2O 0.264 103.50

S8 0.005 2.18

TOTAL 1.0 391.48


As in the previous case all of the produced sulfur is condensed through the condenser and then withdrawn from the collecting pits.




42




0.999 = 2.185 kgmol


H2O going in Stream-11 = 103.50 kgmol – 0.002 kgmol

= 103.48 kgmol




S8 0.999 2.18

H2O 0.001 0.002

TOTAL 1.0 2.20



H2S 0.037 2.98

CO2 0.192 76.16

N2 0.520 205.20

SO2 0.018 1.45

H2O 0.232 103.48

TOTAL 1.0 389.30


H2S = 0.007

CO2 = 0.194

N2 = 0.524

SO2 = 0.003

H2O = 0.264

S8 = 0.005

Flow Rate = ?

S8 = 0.999

H2O = 0.111

Flow Rate = ?

H2S = ?

CO2 = ?

N2 = ?

SO2 = ?

H2O = ?


43




H2S 0.037 2.98

CO2 0.192 76.16

N2 0.520 205.20

SO2 0.018 1.45

H2O 0.232 103.48

TOTAL 1.0 389.30


Fig 3.8: Material balance across reactor R-102





= 2.76 kgmol




H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232

Flow Rate = ?

H2S = ?

CO2 = ?

N2 = ?

SO2 = ?

H2O = ?

S8 = ?


44



= 2.76 kgmol


S8 produced = 3


= 0.51 kgmol




H2S 0.0005 0.21

CO2 0.120 76.16

N2 0.528 205.20

SO2 0.0001 0.072

H2O 0.273 106.27

S8 0.001 0.51

TOTAL 1.0 388.43


As in the previous cases all of the produced sulfur is condensed through the condenser and then withdrawn from the collecting pits.




H2S = 0.005

CO2 = 0.120

N2 = 0.528

SO2 = 0.0001

H2O = 0.273

S8 = 0.001 Flow Rate = ?

S8 = 0.999

H2O = 0.111

Flow Rate = ?

H2S = ?

CO2 = ?

N2 = ?

SO2 = ?

H2O = ?


45


0.999 = 0.52 kgmol


H2O going in Stream-14 = 106.27 kgmol – 0.0005 kgmol

= 106.27 kgmol

Now the Stream-S24and Stream-14 compositions are as follows:


Component Mole fraction Flow rat-FS24 (kgmol/hr)

S8 0.999 0.51

H2O 0.001 0.0005

TOTAL 1.0 0.52



H2S 0.0005 0.21

CO2 0.196 76.16

N2 0.530 205.20

SO2 0.0001 0.072

H2O 0.274 106.27

TOTAL 1.0 387.91

3.11 Final Calculations

Total S8 produced from SRU = S8 withdrawn from condenser E-100 + S8

withdrawn from condenser E-102 + S8

withdrawn from condenser E-104 + S8

withdrawn from condenser E-106

= 3.90 kgmol + 6.37 kgmol + 2.18 kgmol +

0.51 kgmol

= 12.99 kgmol = 80 tons/day

Chapter 4 ENERGY BALANCE OF THE SULFUR RECOVERY UNIT

(SRU)

4.1 Introduction

he energy balance across the proposed sulfur recovery unit (SRU) is done by the conservation equation of energy, as is done conventionally. A system must be defined to account for the streams entering and leaving. In our case the obvious selection is the sulfur recovery unit (SRU) itself while all the other premises are considered surroundings. Some preliminary bases are to be specified for the

sake of convenience in the calculations. Following specifications are taken to meet the above mentioned situation:

Time of operation: 1 hr

Ambient temperature: 25oC

Ambient pressure: 1 atm

Now the energy balance calculations are made by using the following equation for the law of conservation of energy:

Amount of energy Amount of energy Amount of energy

entering the system - leaving the system + generated within the -

through the boundaries through the boundaries system boundaries

Amount of energy Amount of energy

consumed within the = accumulated within the (4.1)

system boundaries system boundaries

Furthermore, all the enthalpies of the streams are calculated by the following relation:

Q = Σ (mCp) ∆T (4.2)

T

Chapter 4 Energy Balance of the SRU

47

Whereas:

Q = amount of heat contained by the stream (kJ/hr)

m = molar flow rate of the stream (kgmol/hr)

Cp = Heat capacity of the stream (kJ/kgmol-oC)

∆T = Temperature of the stream (oC)

The chemical reactions taking place in the process are:

Main Reactions1:

1- H2S + 3/2O2 SO2 + H2O (4.3)

(∆H = -4147.20 kJ/kgmol)

2- SO2 + 2H2S 3/8S8 + 2H2O (4.4)

(∆H = -1165.60 kJ/kgmol)

Side Reactions:

3- CH4 + 2O2 CO2 + 2H2O (4.5)

(∆H = -891.0 kJ/kgmol)

4- C2H6 + 7/2O2 2CO2 + 2H2O (4.6)

(∆H = -1560.0 kJ/kgmol)

5- C3H8 + 5O2 3CO2 + 4H2O (4.7)

(∆H = -2220.0 kJ/kgmol)

NOTE: In all the diagrams the compositions are mentioned in mole fraction basis

4.2 Overall Energy Balance

1 All heat of reaction data is taken from: http://en.wikipedia.org/wiki/Claus_process


48

Fig 4.1: Overall energy balance across SRU

Stream-01 H2S CO2 CH4 H2O C2H6 C3H8 Total

m (kgmol/hr) 104.03 74.83 0.95 18.80 0.16 0.02 198.80

Cp (kJ/kgmol-oC) 35.65 41.29 40.35 34.50 64.44 92.91 -

∆T (oC) 190.6 190.6 190.6 190.6 190.6 190.6 -

Q01 (kJ/hr) 706872 588902 7306 123623 1965 354 1429098

Stream-02 O2 N2 Total

m (kgmol/hr) 54.54 205.20 259.70

Cp (kJ/kgmol-oC) 29.49 29.30 -

∆T (oC) 45.0 45.0 -

Q02 (kJ/hr) 72377 270556 342940

4.3 Energy Balance across Furnace F-100


Temp. = 215.6 o

C

H2S = 0.523

CO2 = 0.376

CH4 = 0.004

H2O = 0.010

C2H6 = 0.0008

C3H8 = 0.0001


Temp. = 70 o

C

O2 = 0.210

N2 = 0.790


Temp. = ?

H2S = 0.116

CO2 = 0.183

N2 = 0.50

SO2 = 0.058

H2O = 0.140

S8 = 0.010


Temp. = 215.6 o

C

H2S = 0.523

CO2 = 0.376

CH4 = 0.004

H2O = 0.010

C2H6 = 0.0008

C3H8 = 0.0001


Temp. = 70 o

C

O2 = 0.210

N2 = 0.790

Flow Rate = 13 kgmol/hr

Temp. = 124.4 o

C

S8 = 0.999

H2O = 0.001

Fig 4.2: Energy balance across furnace F-100


49

Enthalpy of Stream-01 = 1429098.75 kJ/hr

Enthalpy of Stream-02 = 342939.52 kJ/hr

Heat of reactions of all

reactions taking place = (-4147.20 - 1165.60 - 891.0 - 1560.0 - 2220.0) kJ/hr = -9983.80 kJ/hr

in the furnace

Total amount of enthalpy = (1429098.75 + 342939.52 - 9983.80) kJ/hr = 1762054.47 kJ/hr

within the furnace

Now we calculate the output temperature of the Stream-03 using equ. 4.1:

Q = Σ (mCp) ∆T

1762054.47 kJ/hr = 416.0 kgmol/hr × Cp × (Tout-25 oC)

By iteration, the outlet temperature comes out to be 1177 oC (2150 oF)

Stream-03 H2S CO2 N2 SO2 H2O S8 Total

m (kgmol/hr) 48.58 76.16 205.20 24.25 58.0 3.90 416.0

Cp (kJ/kgmol-oC)

51.19 58.36 34.26 57.01 45.69 655.20 -

∆T (oC) 1152.0 1152.0 1152.0 1152.0 1152.0 1152.0 -

Q03 (kJ/hr) 2864805 5120291 8098735 1592631 3052823 2943682 23672970

4.4 Energy Balance across Boiler B-100

Fig 4.3: Energy balance across boiler B-100


Temp. = 1177 o

C

H2S = 0.116

CO2 = 0.183

N2 = 0.50

SO2 = 0.058

H2O = 0.140

S8 = 0.010


Temp. = 649 o

C

H2S = 0.116

CO2 = 0.183

N2 = 0.50

SO2 = 0.058

H2O = 0.140

S8 = 0.010


50

The waste heat recovery boiler extracts such an amount of energy from the stream-03 that the outlet temperature of the stream leaving the boiler; stream-04, becomes equal to 649 oC (1200 oF).


m (kgmol/hr) 48.58 76.16 205.20 24.25 58.0 3.90 416.0

Cp (kJ/kgmol-oC)

44.77 52.72 32.24 53.89 40.43 165.50 -

∆T (oC) 624.0 624.0 624.0 624.0 624.0 624.0 -

Q04 (kJ/hr) 1357154 2505456 4128164 815463 1463242 402760 10672242


Fig 4.4: Energy balance across condenser E-100

The condenser E-100 reduces the temperature of the stream-04 from 649 oC (1200 oF) to 124.4 C (256 oF) which is the dew point temperature of rhombic sulfur.

Stream-05 H2S CO2 N2 SO2 H2O Total

m (kgmol/hr) 48.58 76.16 205.20 24.25 57.90 412.0

Cp (kJ/kgmol-oC) 35.37 40.69 29.56 42.89 79.20 -

∆T (oC) 99.4 99.4 99.4 99.4 99.4 -

Q05 (kJ/hr) 170796 308035 602931 103384 455816 1640964

Stream-S21 S8 H2O Total

m (kgmol/hr) 3.90 0.004 3.904

Cp (kJ/kgmol-oC) 33.75 79.20 -

∆T (oC) 99.4 99.4 -

QS21 (kJ/hr) 13083 31.4 13115


Temp. = 649 o

C

H2S = 0.116

CO2 = 0.183

N2 = 0.50

SO2 = 0.058

H2O = 0.140

S8 = 0.010


Temp. = 124.4 o

C

H2S = 0.117

CO2 = 0.184

N2 = 0.50

SO2 = 0.058

H2O = 0.140 Flow Rate = 3.904 kgmol/hr

Temp. = 124.4 o

C

S8 = 0.999

H2O = 0.111


51

4.6 Energy Balance across Heat Exchanger E-101

Fig 4.5: Energy balance across heat exchanger E-100

The heater heats the incoming stream-05 from 124.4 oC (256 oF) to 248.8 oC (480 oF) which is the required temperature of the first reactor R-100.


m (kgmol/hr) 48.58 76.16 205.20 24.25 57.90 412.0

Cp (kJ/kgmol-oC) 37.27 44.07 30.20 46.67 35.43 -

∆T (oC) 223.8 223.8 223.8 223.8 223.8 -

Q06 (kJ/hr) 405207 751156 1386897 253285 459102 3255648


Fig 4.6: Energy balance across reactor R-100

Enthalpy of Stream-06 = 3255648 kJ/hr


Temp. = 248.8 o

C

H2S = 0.117

CO2 = 0.184

N2 = 0.50

SO2 = 0.058

H2O = 0.140


Temp. = 124.4 o

C

H2S = 0.117

CO2 = 0.184

N2 = 0.50

SO2 = 0.058

H2O = 0.140


Temp. = 248.8 o

C

H2S = 0.117

CO2 = 0.184

N2 = 0.50

SO2 = 0.058

H2O = 0.140


Temp. = ?

H2S = 0.036

CO2 = 0.190

N2 = 0.511

SO2 = 0.018

H2O = 0.228

S8 = 0.015


52

Heat of reaction taking = -1165.60 kJ/hr

place in the reactor

Total amount of enthalpy = 3255648 kJ/hr – 1165.60 kJ/hr = 3254482 kJ/hr

within the furnace


Q = Σ (mCp) ∆T

3254482 kJ/hr = 401.50 kgmol/hr × Cp × (Tout-25 oC)

By iteration, the outlet temperature comes out to be 354.4 oC (670 oF)


m (kgmol/hr) 14.62 76.16 205.20 7.27 91.86 6.36 401.50

Cp (kJ/kgmol-oC)

39.14 46.63 30.72 49.20 36.61 60.32 -

∆T (oC) 329.4 329.4 329.4 329.4 329.4 329.4 -

Q07 (kJ/hr) 188491 1169811 2076453 117821 1107770 126370 4786717




Temp. = 354.4 o

C

H2S = 0.036

CO2 = 0.190

N2 = 0.511

SO2 = 0.018

H2O = 0.228

S8 = 0.015


Temp. = 124.4 o

C

H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232 Flow Rate = 6.38 kgmol/hr

Temp. = 124.4 o

C

S8 = 0.999

H2O = 0.111


53


m (kgmol/hr) 14.62 76.16 205.20 7.27 91.86 395.12

Cp (kJ/kgmol-oC) 35.37 40.69 29.56 42.89 79.20 -

∆T (oC) 99.4 99.4 99.4 99.4 99.4 -

Q08 (kJ/hr) 51400 308035 602931 30994 723166 1716528



The heater heats the incoming stream-08 from 124.4 oC (256 oF) to 204.4 oC (400 oF) which is the required temperature of the second reactor R-101.


m (kgmol/hr) 14.62 76.16 205.20 7.27 91.86 395.12

Cp (kJ/kgmol-oC) 36.45 42.91 29.97 45.52 35.0 -

∆T (oC) 179.4 179.4 179.4 179.4 179.4 -

Q09 (kJ/hr) 95602 586284 1103282 59369 576789 2421325



m (kgmol/hr) 6.37 0.0067 6.38

Cp (kJ/kgmol-oC) 33.75 79.20 -

∆T (oC) 99.4 99.4 -

QS22 (kJ/hr) 21370 52.7 21422


Temp. = 124.4 o

C

H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232


Temp. = 204.4 o

C

H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232


54






within the furnace


Q = Σ (mCp) ∆T




m (kgmol/hr) 2.98 76.16 205.20 1.45 103.50 2.18 391.48

Cp (kJ/kgmol-oC) 37.18 43.93 30.17 46.52 35.38 44.48 -

∆T (oC) 218.3 218.3 218.3 218.3 218.3 218.3 -

Q10 (kJ/hr) 24187 730368 1351470 14725 799377 21167 2941295



Temp. = 204.4 o

C

H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232


Temp. = ?

H2S = 0.007

CO2 = 0.194

N2 = 0.524

SO2 = 0.003

H2O = 0.264

S8 = 0.005


55



m (kgmol/hr) 2.98 76.16 205.20 1.45 103.48 389.30

Cp (kJ/kgmol-oC) 35.37 40.69 29.56 42.89 79.20 -

∆T (oC) 99.4 99.4 99.4 99.4 99.4 -

Q11 (kJ/hr) 10477 308035 602931 6181 814644 1742270




m (kgmol/hr) 2.18 0.002 2.20

Cp (kJ/kgmol-oC) 33.75 79.20 -

∆T (oC) 99.4 99.4 -

QS23 (kJ/hr) 7313 15.7 7329


Temp. = 243.3 o

C

H2S = 0.007

CO2 = 0.194

N2 = 0.524

SO2 = 0.003

H2O = 0.264

S8 = 0.005


Temp. = 124.4 o

C

H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232


Temp. = 124.4 o

C

S8 = 0.999

H2O = 0.111


Temp. = 124.4 o

C

H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232


Temp. = 196.6 o

C

H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232


56

The heater heats the incoming stream-11 from 124.4 oC (256 oF) to 196.6 oC (386 oF) which is the required temperature of the third reactor R-102.


m (kgmol/hr) 2.98 76.16 205.20 1.45 103.48 389.30

Cp (kJ/kgmol-oC) 36.42 42.70 29.93 45.19 34.92 -

∆T (oC) 171.6 171.6 171.6 171.6 171.6 -

Q12 (kJ/hr) 18624 558049 1053904 11244 620080 2261902







within the furnace


Q = Σ (mCp) ∆T




Temp. = 196.6 o

C

H2S = 0.037

CO2 = 0.192

N2 = 0.520

SO2 = 0.018

H2O = 0.232


Temp. = ?

H2S = 0.007

CO2 = 0.194

N2 = 0.524

SO2 = 0.003

H2O = 0.264

S8 = 0.005


57


m (kgmol/hr) 0.21 76.16 205.20 0.072 106.27 0.51 388.43

Cp (kJ/kgmol-oC) 36.59 42.99 29.98 45.50 35.02 40.70 -

∆T (oC) 182.2 182.2 182.2 182.2 182.2 182.2 -

Q13 (kJ/hr) 1400 596544 1120875 597 678071 3782 2401270




m (kgmol/hr) 0.21 76.16 205.20 0.072 106.27 387.91

Cp (kJ/kgmol-oC) 35.37 40.69 29.56 42.89 79.2 -

∆T (oC) 99.4 99.4 99.4 99.4 99.4 -

Q14 (kJ/hr) 738 308035 602931 307 836608 1748621


m (kgmol/hr) 0.51 0.0005 0.52

Cp (kJ/kgmol-oC) 33.75 79.20 -

∆T (oC) 99.4 99.4 -

QS24 (kJ/hr) 1711 4 1714.8


Temp. = 207.2 o

C

H2S = 0.007

CO2 = 0.194

N2 = 0.524

SO2 = 0.003

H2O = 0.264

S8 = 0.005


Temp. = 124.4 o

C

H2S = 0.0005

CO2 = 0.196

N2 = 0.530

SO2 = 0.0001

H2O = 0.274


Temp. = 124.4 o

C

S8 = 0.999

H2O = 0.111

Chapter 5 EQUIPMENTS DESIGN

5.1 Design of Reaction Furnace (F-100)

Fig 5.1: Material and energy flow across furnace F-100

Here the equipment called muffle furnace is actually a plug flow reactor provided with refractory linings to avoid material damage on account of liberation of high exothermic heats of reaction and combustion. In this thermal reactor we treat acid gas with air to produce sulfur dioxide which then reacts in the Claus three stage reactors to produce elemental sulfur. As we are oxidizing or burning hydrogen sulfide in the presence of air, the reaction will be highly exothermic, thus liberating high heat contents, so we named it a furnace. Design is totally on reactor based calculations.

The basic design steps included are (keeping in mind that due to absence of kinetic data the reactor has to be designed based upon the concept of residence time):

STEP-1: Estimation of Residence Time (τ)

STEP-2: Calculation of Volumetric Flow Rate (u)


Temp. = 70 o

C

Pressure = 101300 Pa

O2 = 0.210

N2 = 0.790


Temp. = 215.6 o

C

H2S = 0.523

CO2 = 0.376

CH4 = 0.004

H2O = 0.010

C2H6 = 0.0008

C3H8 = 0.0001


Temp. = 1177 o

C

H2S = 0.116

CO2 = 0.183

N2 = 0.500

SO2 = 0.058

H2O = 0.140

S8 = 0.010

Chapter 5 Equipments Design

59

STEP-3: Calculation of Volume of Reactor (V)

STEP-4: Calculation of Length of Reactor (L)

STEP-5: Selection of Refractory Material

STEP-6: Calculation of Internal Diameter of the Reactor (Di)

STEP-7: Calculation of Thickness of Reactor (t)

STEP-8: Calculation of Outer Diameter (Do)

STEP-9: Calculation of the Volume of Steel Required (Vs)

STEP-10: Calculation of the Mass of Steel Required (ms)

STEP-11: Calculation of Pressure Drop (ΔP)

The assumptions taken before coming to the actual design are:

• Plug flow.

• Steady state.

• Constant density

• Constant tube diameter.

• No mixing in the axial direction.

• Complete mixing in the radial direction.

• A uniform velocity profile across the radius.

STEP-1: Estimation of Residence Time (τ)

Residence time (T) = Volume of reactor (V)

Volumetric flow rate (u)

Or Volume of Reactor (V) = Residence time (T) × Volumetric flow rate (u) (5.1)

The residence time is actually based upon the results obtained from the information and calculations of the kinetics of the reactions involved. Since kinetic data is not available we have to take help from literature. The residence time is taken as 2.5 seconds.1

Residence time (T) = 2.5 s

1 Perry’s Chemical Engineers’ Handbook, Volume-1, 7

th Edition by Robert H. Perry, Don Green


60

STEP-2: Calculation of Volumetric Flow Rate (u)

For the calculation of volumetric flow rate we proceed as:

ε = uo - ui

ui

(5.2)

Whereas:

ε = Expansivity (dimensionless)

uo = Outlet volumetric flow rate (m3/hr)

ui = Inlet volumetric flow rate (m3/hr)

Now:

Inlet volumetric flow rate (ui) = Inlet mass feed flow rate (mi)

Density of the inlet stream (ρi) (5.3)

Inlet mass feed flow rate (mi) = Mass flow rate of stream-01 + Mass flow rate of stream-02

For the calculation of the mass flow rate of both streams we must know the average molecular weight of both the streams.

Stream-01 composition:

Component Mole fraction Molecular weight

(kg/kgmol)

H2S 0.523 34.0

CO2 0.376 44.0

CH4 0.004 16.0

H2O 0.010 18.0

C2H6 0.0008 30.0

C3H8 0.0001 44.0

Average molecular Weight of stream-01 = (0.523 × 34.0) + (0.376 × 44.0) + (0.004 × 16.0) + (0.010

× 18.0) + (0.0008 × 30.0) + (0.0001 × 44.0)

= 34.60 kg/kgmol

Similarly:

Stream-02 composition:

Component Mole fraction Molecular weight

(kg/kgmol)

O2 0.210 32.0

N2 0.790 28.0


61

Average molecular Weight of stream-02 = (0.210 × 32.0) + (0.790 × 28.0)

= 28.84 kg/kgmol

Now:

Mass flow rate of stream-01 =

= 6878.50 kg/hr

Mass flow rate of stream-02 =

= 7490.0 kg/hr

Inlet mass feed flow rate (mi) = 6878.50 kg/hr + 7490.0 kg/hr = 14368.50 kg/hr

Now we are to calculate the density of the inlet stream (Di).

We know that the general equation of gas is given by the following relation:

PV = nRT

Whereas:

P = Pressure of the gas (atm) = 1 atm

V = Volume of the gas (dm3)

n = no. of moles of the gas (kgmol)

R = Universal gas constant (atm-dm3/ K-kgmol) = 0.0821 atm-dm3/ K-kgmol

T = Temperature of the gas (K) = 416.2 K

Or:

PV = M

M.WRT

Whereas:

M = Mass of the gas (kg)

M.W = Molecular weight of the gas (kg/kgmol)

198.80 kgmol 34.60 kg hr 1 kgmol



62

M

V =ρi =

P × M.W

RT

For the calculation of molecular weight of the gas we proceed as follows:

The composition of the stream formed after the mixing of stream-01 and stream-02 is as follows:

Mixed stream composition:

Component Flow rate

(kgmol/hr) Mole

Fraction Molecular Weight

(kg/kgmol)

H2S 104.03 0.226 34.0

CO2 74.83 0.163 44.0

CH4 0.95 0.002 16.0

H2O 18.8 0.041 18.0

C2H6 0.16 0.0003 30.0

C3H8 0.02 0.00004 44.0

O2 54.54 0.119 32.0

N2 205.2 0.447 28.0

Molecular weight of the mixed stream = (0.226 × 34.0) + (0.163 × 44.0) + (0.002 × 16.0) +

(0.041 × 18.0) + (0.0003 × 30.0) + (0.00004 × 44.0)

+ (0.119 × 32.0) + (0.447 × 28.0)

= 32.10 kg/kgmol

Now:

ρi = 1 × 32.10

0.0821 × 416.2 = 0.94 kg/m3

Now from equ. 5.3:

Inlet volumetric flow rate (ui) = 14368.50 kg/hr

0.94 kg/m3 = 15285.60 m3/hr

Now we must calculate the outlet volumetric flow rate (uo) of stream-03:

Outlet volumetric flow rate (uo) = Outlet mass feed flow rate (mo)

Density of the inlet stream (ρo)

(8.4)


63

For the calculation of the mass flow rate of outlet stream we must know the average molecular weight of this stream.


Component Mole fraction Molecular Weight

(kg/kgmol)

H2S 0.116 34.0

CO2 0.183 44.0

N2 0.493 28.0

SO2 0.058 24.25

H2O 0.140 18.0

S8 0.010 256.0

Average molecular weight of the stream-03 = (0.116 × 34.0) + (0.183 × 44.0) + (0.493 ×

28.0) + (0.058 × 24.25) + (0.140 × 18.0) +

(0.010× 256.0)

= 32.28 kg/kgmol

Mass flow rate of stream-03 (mo) =

= 13428.48 kg/hr

Now we are to calculate the density of the outlet stream (Do).

ρo = P × M.W

RT

ρo = 1 × 32.28

0.0821 × 1450 = 0.271 kg/m3

Now from equ. 5.4:

Outlet volumetric flow rate (uo) = 13428.48 kg/hr

0.271 kg/m3 = 49551.50 m3/hr

Now from equ. 5.2:

ε = 49551.50 - 15285.60

15285 .60 = 2.24

Now the volumetric flow rate is given by:



64

u = ui (1 + εXa) (5.5)

u = 14368.50 (1 + 2.24 × 0.44) = 4.8 m3/s

STEP-3: Calculation of Volume of Reactor (V)

The volume of the reactor is calculated by the equ. 5.1:

V = 2.5 s × 4.8 m3/s = 12.0 m3

Volume of the reactor (V) = 12.0 m3

STEP-4: Calculation of Length of Reactor (L)

For high rate of mixing and for reaction completion we take small length to diameter ratio as:

L / Di = 3.375

Whereas:

L = Length of the reactor (m)

Di = Internal diameter of the reactor (m)

L = 3.375 Di

As for the cylindrical shell, the volume is given by:

V = (π

4) Di

2L

V = (π

4) Di

2 (3.375 Di)

V = 3.375 (π

4) Di

3

So:

Di = 4V

π×3.375

3 = 1.65 m

And:

L = 3.375 (1.65) = 5.58 m

Length of the reactor (L) = 5.58 m


65

STEP-5: Selection of Refractory Material

As the reactions taking place in the furnace F-100 are all exothermic and the temperature within the furnace rises to 1177 OC (2150 OF), some kind of refractory material is used to avoid the damage of the material of construction of the plug flow reactor on account of high thermodynamic temperature. The refractory thickness is so adjusted that the allowable temperature for shell material will be 354.5 OC (670 OF) because we have suggested using the SA-516 material for the shell construction. By using 0.08 m (3.5 in) refractory (Casto Last G; as it is more convenient for operating under high temperatures and resistant towards sulfur and its products) the shell internal diameter is now increased.

STEP-6: Calculation of Internal Diameter of the Reactor (Di)

Di = 1.65 + 2 (0.08) = 1.81 m

Internal diameter of the reactor (Di) = 1.81 m

STEP-7: Calculation of Thickness of the Reactor (t)

Design pressure (Pi) = 344.7 kPa (50 psi)

Stress factor at 354.5 OC for SA-516 (f) = 3.08 × 104 kPa (4475 psi)

Corrosion allowance (Cc) = 3.17 × 10-3 (0.125 in)

Joint efficiency (J) = 0.85

Now the thickness of the shell is given by:

t = Pi ×Di

2f ×J + Cc

t = 344.7 × 1.81

2(3.08 × 104) × 0.85 = 0.02 m

Thickness of the reactor (t) = 0.02 m

STEP-8: Calculation of Outer Diameter of the Reactor (Do)

Now the outer diameter (Do) is calculated by:

Do = Di + 2t

Do = 1.81 + 2 (0.02) = 1.85 m

Outer diameter of the reactor (Do) = 1.85 m

STEP-9: Calculation of the Volume of Steel Required (Vs)

Volume of the austenitic steel required is calculated by:

Vs = π

4 [Do

2-Di2] × L = 1.38 m3


66

Volume of steel required (Vs) = 0.64 m3

STEP-10: Calculation of the Mass of Steel Required (ms)

Specific gravity of steel = 7.84

Density of steel = 7938.311 kg/m3

Now the mass of steel is calculated by:

Mass of steel = Density of steel × Volume of steel

Mass of steel = 7938.311 × 0.64 = 5080 kg

Mass of steel required = 5080 kg

STEP-11: Calculation of Pressure Drop (ΔP)

Pressure drop occurs only if there is a catalyst bed and is calculated by the Ergun equation. Since no catalyst bed is not used over here so ΔP = 0, since a plug flow reactor operating without any catalyst has 0 pressure drop.1

SPECIFICATION SHEET OF FURNACE (F-100)

IDENTIFICATION Plug Flow Reactor (provided with refractory lining)

NUMBER 1

FUNCTION Conversion of H2S into SO2

OPERATION Continues

TEMPERATURE 1177 OC

OPERATING PRESSURE 450.0 kPa

DESIGN PRESSURE 900.0 kPa

LENGTH 6.10 m

OUTER DIAMETER 1.85 m

MATERIAL OF CONSTRUCTION Austenitic Steel, ASTM Code: SA-516-17

CAPACITY 12.0 m3

MASS OF THE STEEL 5080 kg

REFRACTORY Fireclay

1 A. Kayoed Coker, “Modeling of chemical kinetics and reactor design”.


67

5.2 Design of Waste Heat Recovery Boiler (B-100)

Fig 5.2: Material and energy flow across waste heat recovery boiler B-100

The Design steps for the design calculations are as follows:

STEP-1: Calculation of Mass Flow Rate of Cooling Water (mB100)

STEP-2: Calculation of LMTD (ΔTLMTD)

STEP-3: Assuming a Value of Ud

STEP-4: Calculation of Total Heat Transfer Area (At)

STEP-5: Tube Specifications

STEP-6: Correction of Area (A)

STEP-7: Correction of Ud

STEP-8: Shell Side Heat Transfer Coefficient (ho) Calculations

STEP-9: Tube Side Heat Transfer Coefficient (hi and hi,o) Calculations

STEP-10: Overall Clean Heat Transfer Coefficient (Uc) Calculations

STEP-11: Calculation of Dirt Factor (Rd)

STEP-12: Calculation of Shell Side Pressure Drop (ΔPs)

STEP-13: Calculation of Tube Side Pressure Drop (ΔPT)

Now each design step is put into practical approach as follows:


Temp. = 1177 o

C

H2S = 0.116

CO2 = 0.183

N2 = 0.500

H2O = 0.140

SO2 = 0.058

S8 = 0.010


Temp. = 1177 o

C

H2S = 0.116

CO2 = 0.183

N2 = 0.500

SO2 = 0.058

H2O = 0.140

S8 = 0.010


68

STEP-1: Calculation of Mass Flow Rate of Cooling Water (mB100)

Mass flow rate of cooling water is already calculated in energy balance which comes out to be equal to 3.24 x 103 kg/hr

STEP-2: Calculation of LMTD (ΔTLMTD)

Hot fluid (stream-03) inlet temperature = 1177 oC

Hot fluid (stream-03) outlet temperature = 649 oC

Cold fluid (cooling water) inlet temperature = 25 oC

Super heated steam outlet temperature = 440 oC

(ΔTLMTD)preheat = 585 oC

(ΔTLMTD)superheat = 659 oC

(ΔTLMTD)weighted = Qtotal

(Q/LMTD)preheat + (Q/LMTD)superheat

For Qpreheat and Qsuperheating we proceed as follows:

Qpreheat = mB100Cp (100-25) = 1.02 x 106 kJ/hr

Qsuperheating = mB100Cp (440-100) + λ = 9.30 x 106 kJ/hr

Qtotal = 1.03 x 107 kJ/hr

Whereas:

λ = latent heat of vaporization of water = 1.87 kJ/hr

Now:

(ΔTLMTD)weighted = 649 oC


• Range = 10 – 250 W/m2-K

• Selected Ud = 140 W/m2-K

(depending upon hot and cold fluid)


69

STEP-4: Calculation of Total Heat Transfer Area (At)

At = Q/(Ud x LMTD) = 31.4 m2

STEP-5: Tube Specifications

• Length:

Commonly used = 3.6 m, 4.2 m, 4.8 m

Selected length (L) = 3.6 m

• Tube Outside Diameter = 0.0254 m

• Pitch = 0.032 m (triangular)

• BWG = 16

• Wall thickness = 0.0016 m

• Tube inside diameter (da) = 0.022 m

• Flow area (at) = 0.00038 m2

Outside surface per linear ft = 0.079 m

• Area of one tube = π x OD x L = 0.28 m2

• No of tubes (NT):

By formula = At / Area of one tube = 109

Nearest count = 91

• Correspondingly shell inside diameter:

0.38 m

40% extra diameter for overdesign and vapors is added:

New shell ID = 0.5 m

• Bundle diameter (db):

NT = at (db/da)2.29

db = 0.3 m

• Clearance (C) = Shell ID – db = 0.08 m


70

• Baffle spacing (B):

Range = 0.2 to 1 times shell ID

Calculated = 0.8 (0.38) = 0.3 m

No. of baffles = Length of tubes = 12

Baffle spacing

STEP-6: Correction of Area (A)

A = Nt x (tube outside surface per linear ft) x (tube length) = 25.90 m2

STEP-7: Correction of Ud

Ud = Q / (A x LMTD) = 170 W/m2-K

STEP-8: Shell Side Heat Transfer Coefficient (ho) Calculations

PREHEATING:

Reynolds No. (Re) = (G x De) / μ

For 0.0254 m tube OD and 0.032 m triangular pitch:

De = 0.02 m

Re = 191.0

jH = 7.0 2

ho = jH (k / De) (Cp /k)1/3 = 3706 W/m2-K

SUPERHEATING

Flow area (as) = (Shell ID x C x B) / Pt = 0.285 m2

Mass velocity (Gs) = mB100 / as = 11.3 x 103 kg/hr-m2.

Reynolds No. (Re) = 6532

jH = 42 2

ho = jH (k / De) (Cp /k)1/3 = 752 W/m2-K


71

STEP-9: Tube Side Heat Transfer Coefficient (hi and hi,o) Calculations

Flow area (aT) = 0.034 m2

Mass Velocity (Gt) = mt/aT = 4.4 x 105 kg/m2-hr

Reynolds No. (Re) = 2.05 x 104

jH = 90

k = 0.0084 W/m-K 2

hi = 0.023 (k/D) (DiGt/μ)0.8 (Cp μ/k)0.33 = 342.0 W/m2-K

hi,o = (hi x ID)/OD = 296.0 W/m2-K

STEP-10: Overall Clean Heat Transfer Coefficient (Uc) Calculations

PREHEATING

Up = (hi,o x ho,preheat)/(hi,o + ho,preheat) = 274.0 W/m2-K

Ap = clean area for preheating = Qpreheat / (Uc x ΔTp) = 1.76 m2

SUPERHEAT

Uv = (hi,o x ho,superheat) / (hi,o + ho,superheat) = 212 W/m2-K

Av = clean area for vaporization = Qsuperheat / (Uv x ΔTs) = 18.4 m2

Ac = total clean area = Ap + Av = 20.16 m2

Uc = [(Qp/ΔTp) + (Qv/ΔTv)] / Ac = 218 W/m2-K

STEP-11: Calculation of Dirt Factor (Rd)

Rd = (1/Ud) – (1/Uc )= 1.30 x 10-3 m2-K/W

STEP-12: Calculation of Shell Side Pressure Drop (ΔPs)

PREHEATING

Re = 191

f = 0.0048 1


72

Effective length of preheating zone (L’) = L( Ap/Ac) = 0.31 m

No. of crosses = 1 + No. of baffles = 13

ΔPs,preheat = [f x Gs2 x Ds x (N+1)]/(5.022 x 1010 x De x S x φs) = 16.0 kPa

SUPERHEATING

Re = 6532

f = 0.0022

ΔPs,superheat = [f x Gs2 x Ds x (N+1)]/(5.022 x 1010 x De x S x φs) = 29.0 kPa

ΔPs = Total shell side pressure drop = 45.0 kPa

Which is in the allowable limit i-e: 68.0 kPa

STEP-13: Calculation of Tube Side Pressure Drop (ΔPT)

Re = 2.04 x 104

f = 0.001

ΔPt = (f x Gt2 x n x L) / (5.22 x 1010 x Di x S x φ) = 18.0 kPa

ΔPr = Return pressure drop = (4n/S) x (v2/2gc)

v2/2gc = 2.92 kPa 2 (for Gt = 4.4 x 105)

ΔPr = 24 kPa

Total pressure drop on tube side (ΔPT) = 42.0 kPa; Which is in the allowable limit i-e: 68.0 kPa

SPECIFICATION SHEET OF BOILER (B-100)

UNIT Waste Heat Recovery Boiler

TYPE Fire Tube

OBJECTIVE Heat recovery from flue gases

LAYOUT Horizontal

NO OF UNITS 1

SHELL SIDE FLUID Water

TUBE SIDE FLUID Sour Gases

HEAT LOAD 1.03 10 7 kJ/hr

MASS FLOW RATE ON SHELL SIDE 3.24 10 3 kgmol/hr

MASS FLOW RATE ON TUBE SIDE 416.0 kgmol/hr

INLET WATER TEMPERATURE 25 C


73

OUTLET WATER TEMPERATURE 440 C

INLET GASES TEMPERATURE 1177 C

OUTLET GASES TEMPERATURE 649 C

WEIGHTED LMTD 649 C

HEAT TRANSFER AREA 25,9 m2

INSIDE SHELL DIA 0.95 m

NO OF TUBES 91

INSIDE TUBE DIA 0.022 m

OUTSIDE TUBE DIA 0.0245 m

TUBE LENGTH 3.6 m

TUBE PITCH Triangular (0.032 m)

CLEARANCE 0.08 m

SHELL SIDE PRESSURE DROP 45 kPa

TUBE SIDE PRESSURE DROP 42 kPa

5.3 Design of Reactors (R-100, R-101, R-103)

REACTOR R-100:

Fig 5.3: Material and energy flow across reactor R-100

Temp of entering of stream = 522 K Pressure of entering stream = 1.4 atm Flow rate of entering stream =412 kgmol/hr CALC. OF AVG DENSITY: Molar Density = P/RT (R=0.0802 atm-m3/kgmol-k)

= 1.4 / (0.0802)(522) = 0.0327 kgmol/m3


Temp. = 248.8 o

C

H2S = 0.117

CO2 = 0.184

H2O = 0.410

SO2 = 0.058

N2 = 0.500


Temp. = 354.4 o

C

H2S = 0.036

CO2 = 0.196

H2O = 0.228

SO2 = 0.018

N2 = 0.511

S8 = 0.015


74

Avg. density of stream (p) = 1.056 kg/m3 CALC OF VOL FLOWRATE: Mass flow rate of entering stream = Molar flow rate of entering stream/density

= 3.697 kg/sec Vol flow rate = m/p v = 3.50 m3/sec For the conversion of 70% and for the inlet temp of 500 K, space time comes to be 6 sec. Space velocity would be sec = 600/hr CALC OF BED VOLUME: vB = vol. Flow rate/space velocity vB = 21 m3 Void fraction for volume of catalyst bed is 0.4, so that vol. of bed is vB = (1+E)Vb Total vol. Of catalyst bed = vB = 29.4m3 DIAMETER OF BED: Now suppose L/D = 3 So diameter of bed can be calculated as; DB = (4Vb/3*)1/3 = 2.32 m LENGTH OF BED: LB = 3D = 6.96m Provide 10% free space over bed (top and bottoms for support of catalyst bed) so; LENGTH OF REACTOR: LB(1.1) = 7.656m DIAMETER OF REACTOR: As L/D = 3, so diameter of reactor = 2.552 m Vol. OF REACTOR: V = 39.14 m3

CALCULATION OF PRESSURE DROP: Pressure drop in fixed bed reactor can be calculated by using Ergun’s Equation which is given as follows: ∆P/L = {150є2µVs/(1- є)3Dp

2 + 1.75 єpVs2/(1- є)3Dp}

Where Є = Void Fraction Vs = superficial gas velocity; Vs = v/AB Dp= Diameter of catalyst pellet L = Depth of catalyst bed µ = Avg. Viscosity of gas stream


75

p = Avg. Density of gas stream ∆P = {150(0.4)2(0.023)(0.828)/(0.6)3(.003)2 + 1.75(0.4)(1.056)(0.828)2/(0.6)3(.003)}L = {3674 + 766} (6.96) = 13440 N/m2 So pressure drop comes to be comes to be almost .132 atm which is 8.40% of operating pressure. REACTOR WALL THICKNESS: t = (PiDi)/(4fj – 1.2Pi) - Cc Where Pi = Design Pressure = 50 psig Di = Diameter of reactor f = Stress factor for SA-516-17 at 522K is 30845535 N/m2 j = joint efficiency is 0.85 Cc = Corrosion allowance is 2mm t = 50(93.36)/{4(4475)(.125) - 1.2(50)} = 0.05m t = 0.05 m VOLUME OF AUSTENITIC STEEL, ASTM CODE SA-516-17: Do = Di + 2t =2.65 m Volume of steel = π (Do – Di)2L + 2(π/4)Do2t = 1.506 m3 Density of steel = 6294Kg/m3 Mass of steel required = 9479 kg CALCULATION OF Wt. OF CATALYST: Catalyst used is activated Alumina (Al2O3) in form of spherical pellets. Density of catalyst = 967 kg/m3 Weight of catalyst = vol. of catalyst x Density of catalyst WEIGHT OF CATALYST = 28429 kg SPECIFICATION SHEET OF REACTOR 100:

SPECIFICATION SHEET OF REATOR R-100

REACTOR CATAGORY Fixed Bed Reactor

CATALYST ACTIVATED ALUMINA (Spherical Columns)

MODE OF OPERATION Adiabatic Process

SPACE TIME 6sec

VOLUME OF CATALYST BED 29.4m3


76

VOID FRACTION 0.4

VOLUME OF REACTOR 39.14m3

REACTOR LENGTH 7.66m

REACTOR DIAMETER 2.55m

REACTOR PRESSURE DROP 13440N/M2

REACTOR WALL THICKNESS 0.05m

WEIGHT OF CATALYST 28429 kg

WEIGHT OF STEEL 9479 kg

REACTOR R-101: Temp of entering of stream = 477 K Pressure of entering stream = 1.3 atm Flow rate of entering stream =395.12 kgmol/hr CALC. OF AVG DENSITY: Molar Density = P/RT (R=0.0802 atm-m3/kgmol-k) =1.3/(0.0802)(477k) =0.0332 kg-mol/m3 Avg. density of stream p=1.073 kg/m3 CALC OF VOL FLOWRATE: Mass flow rate of entering stream =(Molar flow rate of entering stream)(avg. Mol. Wt) =3.55 kg/sec Vol flow rate = m/p V=3.30 m3/sec For the conversion of 70% and for the inlet temp of 500k, space time comes to be 6sec. Space velocity would be sec=600/hr CALC OF BED VOLUME: vB = vol. Flowrate/space velocity vB = 19.8m3 Void fraction for volume of catalyst bed is 0.4, so that vol. Of bed is vB = (1+E)Vb


77

Total vol. Of catalyst bed = vB = 27.72m3 DIAMETER OF BED: Now suppose L/D = 3 So diameter of bed can be calculated as; DB = (4Vb/3*)1/3 = 2.274m LENGTH OF BED: LB = 3D = 6.82m Provide 10% free space over bed(top and bottoms for support of catalyst bed) so; LENGTH OF REACTOR: LB(1.1) = 7.50m DIAMETER OF REACTOR: As L/D = 3, so diameter of reactor = 2.50m Vol. OF REACTOR: V = 36.86m3 CALCULATION OF PRESSURE DROP: Dynamic pressure drop across a granular bed can be estimated by means of Ergun’s Equation. The equation is dimensionally consistant.

SPECIFICATION SHEET OF REACTOR R-101

REACTOR CATAGORY FIXED BED REACTOR

CATALYST ACTIVATED ALUMINA SPHERICAL PELLETS

MODE OF OPERATION ADIABATIC PROCESS

SPACE TIME 5sec

VOLUME OF CATALYST BED 27.9m3

VOID FRACTION 0.4

VOLUME OF REACTOR 35.64m3

REACTOR LENGTH 6.55m

REACTOR DIAMETER 2.43m

REACTOR PRESSURE DROP 11460N/m2

REACTOR WALL THICKNESS 0.05m

WEIGHT OF CATALYST 24792kg

REACTOR R-102: Temp of entering of stream = 470K Pressure of entering stream = 1.2 atm Flow rate of entering stream =389.30 kg-mol/hr


78

CALC. OF AVG DENSITY: Molar Density = P/RT (R=o.o802 atm-m3/kgmol-k) =1.2/(0.0802)(470k) =0.0318 kg-mol/m3 Avg. density of stream p=1.03 kg/m3 CALC OF VOL FLOWRATE: Mass flow rate of entering stream =(Molar flow rate of entering stream)(avg. Mol. Wt) =3.49 kg/sec Vol flow rate = m/p V=3.38 m3/sec For the conversion of 70% and for the inlet temp of 500k, space time comes to be 6sec. Space velocity would be sec=600/hr CALC OF BED VOLUME: vB = vol. Flowrate/space velocity vB = 20.33m3 Void fraction for volume of catalyst bed is 0.4, so that vol. Of bed is vB = (1+E)Vb Total vol. Of catalyst bed = vB = 28.46m3 DIAMETER OF BED: Now suppose L/D = 3 So diameter of bed can be calculated as; DB = (4Vb/3*)1/3 = 2.294m LENGTH OF BED: LB = 3D = 6.88m Provide 10% free space over bed(top and bottoms for support of catalyst bed) so; LENGTH OF REACTOR: LB(1.1) = 7.57m DIAMETER OF REACTOR: As L/D = 3, so diameter of reactor = 2.524m Vol. OF REACTOR: V = 37.86m3 CALCULATION OF PRESSURE DROP:


79

Dynamic pressure drop across a granular bed can be estimated by means of Ergun’s Equation. The equation is dimensionally consistent.

5.4 Design of Condenser (E-106)

Condenser E-106

Fig 5.4: Material and energy flow across condenser E-106

Following design steps are followed for the condenser E-106: 1) Calculation of Heat duty (From Energy Balance) 2) Calculation of LMTD 3) Assuming Overall Design Coefficient to find out estimated area 4) Assumption of tube length, tube ID etc. 5) Calculation of Reynolds Number(tube) and Tube-side heat transfer coefficient 6) Calculation of Reynolds Number(Shell side) and Shell-side heat transfer coefficient 7) Calculation of tube side and shell side Pressure drops 8) Iteration for best results.

Now each step is practically applied as:

1) Heat duty

Heat duty = 1,644,838 kJ/hr

Flow Rate= 391.48 kgmol/hr

Temp.= 243.3 oC

H2S = 0.007

CO2 = 0.194

N2 = 0.524

SO2 = 0.003

H2O = 0.264

S8 = 0.005


Temp.= 124.4 oC

S8= 0.999

H2O= 0.111


Temp.= 124.4 oC

H2S = 0.0005

CO2 = 0.196

N2 = 0.530

SO2 = 0.0001

H2O = 0.274


80

= 1,559,010 Btu/hr (Process Gas)

2) LMTD Process gas Inlet temperature = 470 °F

Process gas Outlet temperature = 256 °F

Inlet Temperature of Steam = 245 °F

Outlet Temperature of Steam = 289.5 °F

LMTD = (∆T1 - ∆T2)/ ln (∆T1/∆T2)

= 60.6 °F

LMTD Correction Factor = 0.9

∆ T= Corrected LMTD = 54.54 F

3) Overall Design Coefficient (UD) For Gases as hot fluid and Water (steam) as cold fluid

UD = 26 Btu/ft2-oF (2-52) (Table #8 Kern)

4) Tube Side Specifications (Iterated for best results) OD = 1.25 inch = 0.104 ft BWG =11

Thickness = 0.120 in = 0.01 ft

ID = 1.01 in = 0.0842 ft

Flow Area per Tube = s = 0.0055 ft2

Length of tube = 20 ft

Pitch = Triangular =1.25(do) =1.25 x 0.104

PT = 0.13 ft


81

Tube Side Calculations

As Area = π× do× L

Length of Tube = Area

π× do

= 1,099/ π× 0.104

= 3365 ft

Number of tubes = Total Length / Length of one Tube

= 3365/20

= 169

Nearest # of Tubes from literature = 170

Now, to find mass velocity in tubes

Gt =Mass flow rate / (Flow Area per total length of tube)

= 11,152/ (0.8611)

= 12,949 lb/ft2. hr

We find Tube side Reynolds Number

NRe = (inside diameter of tube ×mass velocity)/ viscosity

= Average Viscosity of gases = 0.0393 lb/ft. hr


82

NRe = Di x Gt /

= 27,715

Finding Tube side heat transfer Coefficient

hc = 1.51 x [ 4 Ѓn /μL ] 0.33 [K2

L ρ2

L.g/μL]3

(Eq. 12.4 Kern)

Ѓn= Condensate Loading = W/ 0.5 L Nt (Condensation occurs over 8 ft)

ρL = Density of Liquid (condensate) = 1350 kg/ m3 = 84 lb/ft3

μL = Viscosity of Liquid (condensate) = 225.5 kg/m.sec = 151 lb/ft.sec

KL = Thermal Conductivity of Condensate = 0.146 J/m.sec

hc = 12.8243 Btu/ ft2. hr.F

5) Shell side Calculations

Inside Diameter of shell = 2.083 ft (Kern Pg.#842)

Baffle Spacing = 3.3 ft

Baffle thickness = Tube thickness×2.5

=0.025 ft

Tube Clearance = 0.333× Tube OD

= 0.333×0.105

= 0.0347 ft (Plant Design by Peters & Timmerhaus Pg.# 705)

Now, to find mass velocity in the shell-side

Flow Area for shell side = inside cross-sectional area of shell - outside cross sectional


83

Area of all tubes

(Cross-sectional area = π D2)

4

= π (2.083)2 – 169 (π x [0.105]2)

4 4

Flow Area for shell side = 2.122 ft2

Calculating equivalent Diameter of Shell

Heated Perimeter = 169 x π x OD

= 49.41 ft

Hydraulic Radius (rH) = free area/ heated perimeter

= 0.043 ft

Equivalent Diameter (De) = 4 x rH

= 0.17 ft

Mass Velocity = Mass Flowrate / Flow area

= 35,658 lb/hr. ft2

μ (viscosity) = 0.578 lb.hr/ft

No. of Baffles =Length of tube/ (baffle spacing + baffle thickness)

= 20/(3.3+0.025)


84

= 6

Baffle Height = 0.75×internal shell diameter

= 0.75 x 2.083

= 1.56 ft

According to equation 14-30 of Plant Design by Peter; Timmerhaus

ho= jH x k/De x (CPµ/k)0.33(µ/µw)0.14

jH = 45 (fig. 28 D.Q.Kern)

De =Equivalent Diameter= 0.17 ft

Gs =Mass velocity of shell side = 35,658

CP = 0.4632 Btu/lb. F

µ = 0.578 lb.hr/ft

k = 0.0165 Btu/ hr. ft2 (F/ft) Table 5 Kern

ho= 11 Btu/ ft2.hr.F.

Now, we need to calculate Overall Design Coefficient to check our earlier assumption

As,

(1/Ud) = (1/ho) + (1/hi) (OD/ID) + xw/kμ (OD/ID + xw) 1/hdo +1/hdi (OD/ID)

(Eq. 11-2 Perry)

hdo = 500 (For Gas from Kern Table 12)

hdi = 1000 (For BFW from Kern Table 12)

xw = thickness of wall = 0.01 ft

k = thermal conductivity of condensate = 0.146 J/m.sec


85

Result comes out to be

Ud = 26.7 (assumption is valid)

6) Pressure Drop

Along Tube Side

ΔPi = 0.5( fi x Gi2 x L )/(5.22 x1010 x D x s x Фi) (D.Q. Kern)

fi = 0.0025 (Fig. 26 Kern)

Gi = Mass Velocity in tubes = 12,949 lb/ft2. hr

L = Length of tube= 20 ft

D = OD of tubes= 0.7476 ft

S= Specific gravity of gases= 0.00204

Putting these in above we get the Pressure Drop

ΔPi = 2.8 psi

Along Shell Side

ΔPs =ƒGs2Ds(N+1)/5.22 x105sDe


86

N+ 1 = 12 L /B = 60

Re = 12,051

ƒ = 0.003 ft2/in2

Ds = 2.083 ft

Gs = 35.658 lb/hr. ft2

De = 0.171 ft

Putting these values in above we get shell-side pressure drop

ΔPs = 3.5 psi

SPECIFICATION SHEET OF Condenser (E-106)

Type of Condenser In-tube (Shell and tube)

Shell passes 1

Tube Pattern Triangular

Tube passes 1

Heat Transfer Area 89.5 m2

Number of tubes 170

Type of Baffle Segmental 25% cut

Number of baffles 6

Inlet Temperature of Steam 118.3 °C

Outlet Temperature of Steam 143 °C

Inlet Temperature of organic mixture 243.3 °C

Outlet Temperature of Organic Mixture 124.4 °C


87

Pressure drop at shell side 24 kPa

Tube Side Pressure Drop 19 kPa

Tube side heat transfer coefficient 12.824 Btu/ ft2 F.hr

Shell side heat transfer coefficient(ho) 11 Btu/ ft2 F.hr

Overall heat transfer coefficient(Ud) 26.7 W/ m2-K

8.5 Design of Process Stream Heater (E-105)

Fig 8.5: Material and energy flow across process stream heater E-105

The purpose of the heater is to heat the incoming stream-11 from 124.4 C to 196.9 C as stream-

12. The overall design steps are as follows:

STEP-1: Calculation of Heat Duty.

STEP-2: Calculation of Mass Flow rate.

STEP-3: Assuming a Value of Ud.


88

STEP-4: LMTD Calculations.

STEP-5: Area Calculations.

STEP-6: Clean Coefficient Calculation Uc.

STEP-7: Correction of Ud.

STEP-8: Correction of Area.

STEP-9: Length Calculations.

STEP-10: Calculation of no. of Hairpins.

STEP-11: Calculation of Pressure Drop.

The inner pipe design steps which are to be used are:

• Inner Pipe Flow area

• Mass velocity

• Reynolds Number.

• Prandtl Number.

• Wall Temperature

• Wall Viscosity

• Determination of JH

• Calculation of Inside Heat transfer coefficient

• Pressure Drop Calculations

Similarly, the outer pipe design steps are:

• Annulus Flow area

• Mass velocity

• Equivalent Diameter

• Reynolds Number.

• Prandtl Number.

• Wall Temperature.


89

• Wall Viscosity.

• Determination of JH.

• Calculation of Outside Heat transfer coefficient.

• Pressure Drop Calculations.

The physical properties of the process stream and the steam used are summarized in the following

table for a quick and easy review:

Process Stream Steam

PROPERTY VALUE PROPERTY VALUE

Molecular weight 29.62 kg/kgmol Molecular weight 18.02 kg/kgmol

Heat Capacity (Cp) 34.50 kJ/kgmol-oC Heat Capacity (Cp) 40.06 kJ/kgmol-

oC

Molar Density 0.252 kgmol/m3 Molar Density 0.56 kgmol/m

3

Viscosity 0.73 kg/m-hr Viscosity 0.95 kg/m-hr

Thermal

Conductivity 0.1 kJ/hr-m-

oC

Thermal

Conductivity 0.2 kJ/hr-m-

oC

STEP-1: Calculation of Heat Duty (Q)

Q = mcCpc(t2 –t1)

= 9.6 x 105 kJ/hr

STEP-2: Calculation of Mass Flow Rate (mh)

mh = Q/λ

As:

Qh = Qc

= 30 kgmol/hr


Assuming: Ud = 140 J/s.K.m2


90

=504.0 KJ/hr.m2.oC (For Gases)

STEP-4: LMTD Calculations

LMTD= {∆T1 - ∆T2 }/ ln (∆T1/∆T2)

= 280oC

STEP-5: Area Calculations

Heat Transfer Area = Q/(Ud x LMTD) = 6.87 m2

OUTER PIPE DIMENSIONS.2

IPS = 2 in = 0.0508 m Effective length = 3.65m

Schedule no = 40 Di = 0.053 m

Do = 0.06 m Thickness =0 .007 m

INNER PIPE DIMENSIONS

IPS= 1.25 in = 0.031 m Effective length = 3.65

Schedule no = 40 di = 0.035 m

do = 0.042m Thickness = 0 .007m

Flow Area Ai = π (Di )

2 /4 = 9.6×10-4 m2

Mass Velocity Gi = mi /Ai = 3.12 x 104 Kgmol/m2. hr Reynolds's Number Re = di Gi /µ = 2.0 x 104 (Flow regime is turbulent.)

Prandtl Number

2 D. Q. Kern, “Process Heat Transfer”. (pg 110), Table 11 (pg 844)


91

Pr = Cp µ/k = 10.5

Wall Temperature

∆T across Steam Film = 0.4 (440-160.5) = 111.8 oC

Wall Temp from Steam Side = 272.3 oC

Wall Viscosity

µw of Steam = 0.82 Kg/m.hr

Heat Transfer Coefficient

JHi = 70. 1

JHi = (hidi/ki) (Pr)-0.34 (µ /µ w)-0.14

hi = 91 kJ/hr-m2-oC

Flow Area Ai = π (Di

2 – do2

) /4 = 8.2×10-4 m2

Mass Velocity Go = mo/Ao = 47.5 x 104 Kgmol/m2-hr Equivalent Diameter Deq = 4rH = (Di

2 – do2 )/do = 0.025 m

Reynolds's Number Re = DeqGo/µ = 48 x 104 (Flow regime is turbulent)

Prandtl Number Pr = Cpµ/k = 8.5

Wall Temperature

∆T across Process Stream Film = 0.5 (440-160.5)

= 139.8 oC

Wall Temperature from Gases side = 251 oC


92

Wall Viscosity

µw of Gases = 0.94 Kg/m-hr. 1

Heat Transfer Coefficient

JHo = 165. 2

JHo = (hoDeq/ko) (Pr)-0.34 (µ /µ w)-0.14

ho = 1275 kJ/hr-m2-oC

STEP-6: Clean Coefficient Calculation Uc

1/UC = 1/hi +Di [ ln(Do/Di)]/2KW+Di/ho Do

Uc = 51519 kJ /m2-hr-oC

STEP-7: Corrected Ud

Rd = 1/Ud – 1/Uc = 0.0004 m2-oC-hr/kJ

Corrected Ud = 402 kJ/m2-hr-oC

STEP-8: Corrected Area

Corrected Heat Transfer Area= Q/(Ud x LMTD) = 8.60 m2

STEP-9: Length Calculations

Length Required=A/(π x do) = 8.6/(3.14 x 0.042) = 65.0 m

STEP-10: Calculation of Hairpins Required

Number of hairpins = 10.66

So nearly:

Number of hairpins are 11


93

STEP-11: Pressure Drop Calculations

ANNULUS:

Re = 48 x 104

f = 0.0035 + 0.264 / (Re)0.42 = 0.00046

ΔPs = 4fGo2 L / 2g ρ2 Deq= 6690.2 m

As: V = Go/ρ=18.8 x 104 m/hr

ΔPr = (v2/2g) = 0.0074 m

ΔPt= (ΔPr + ΔPs) x ρ = (6690.2 + 0.0074) x 0.252 x 29.64 = 49971.0 kg/m2 = 49.9 KPa

INNER PIPE:

Friction Factor (f) = 0.0035 + 0.264 = 0.007

(Re)0.42

ΔPi = 4fGi2 L

2g ρ2 di

ΔP = 1.7 x 109/2.7 x 106 = 629.6 m

ΔPi = (ΔPi) x ρ = 629.6 x 0.56 x 18.02 = 6.3 kPa

SPECIFICATION SHEET OF HEATER (E-105)

Identification: Item: Double pipe heat exchanger (E-105)

Function: To heat the process stream from 124.4 0C to 196.6

0C.

HEAT DUTY Heat duty (Q) = 9.6 x 10

5 kJ/hr

A C = 8.6 m2

Ud = 402 kJ/hr-m2-

0C

Inner pipe side: Annulus side:

Fluid handled: Steam Flow rate: 30 kgmol/hr Temperature: 440

0C to 440

0C

Schedule number: 40 Shell material: Carbon Graphite Temperature limit: 500

0C

Pressure drop : 6.3 kPa

Fluid handled: Process gases Flow rate: 389.3 kgmol/hr Temperature: 124.4

0C to 196.6

0C

Schedule number: 40 IPS: 1.25 in = 0.031 m Tube material: Low carbon Steel Temperature limit: 0 - 230

0C

Pressure drop: 49.9 kPa

Chapter 6 PROCESS INSTRUMENTATION & CONTROL

6.1 Introduction

ell-designed and properly applied instrumentation is required for efficient and reliable SRU operation. An SRU with the best designed piping and equipments cannot achieve its intended function without proper instrumentation. Today, extremely good instrumentation is available, and no plant should be limited by lack of proper instrumentation. However,

even with excellent equipment available, instrumentation must be available and must be properly designed and installed. There will always be some residual H2S and SO2 in the SRU area. Instruments must be protected from these corrosive components through purging and sealing. Copper and brass components (tubing, valves, electrical connection, etc.) should never be used in SRU's unless they are inside purged enclosure. Some key instrumentation areas are discussed below.

6.2 General Discussion on Instrumentation of the SRU

Following is a brief discussion about various safety equipments installed in the SRU.

6.2.1 Feed Flow Measurement and Control

Proper feed flow measurement and control is critical to reliable efficient operation of SRU. It is preferred to control the acid gas header and knockout drum pressure inside the SRU battery limits with a flow controller on cascade control from the pressure controller. The feed flow meters should be located upstream of the control valve. This insures constant pressure on the flow meter and avoids the need for pressure compensation of the measured flow. Conventional orifice plates or low-pressure drop venturi meters are normally used.

6.2.2 Combustion Air Control

Control of combustion air feed to the SRU is the most critical for efficient, reliable operation of the SRU. It is preferred to use a combustion air control system that is split into two sections, a main air flow loop based on acid gas flow rate and a trim air flow loop based on the gas tail analyzer. Air demand is calculated from the acid gas flows and used as a feed forward ratio set point for the main air control loop. The main air loop supplies about 90% of the total air to the burner. The trim air loop operates on feed back control from SO2 in the tail gas. The analyzer controller provides a remote set point signal to the trim air loop

W

Chapter 6 Process Instrumentation & Control

95

based on the relationship 2SO2 - H2S = 0. When this relationship satisfied, the optimum amount of combustion air is being supplied to the SRU. If the result is positive, too much air is being fed and the rate should be reduced. Likewise, if the result is negative, too little air is being fed and the rate should be increased. It is apparent that the tail gas analyzer must work properly to achieve optimum air control with efficient, reliable operation. If the air controls are off, inadequate impurity destruction with equipment plugging and/or equipment corrosion may occur. The tail gas analyzer must be properly located, use proper sampling (preferably supplied by the analyzer manufacturer), and be properly celebrated. This requires correct initial design and regular maintenance.

6.2.3 Main Burner and Reaction Furnace

The burner and reaction furnace require a significant amount of instrumentation; however, most is not directly associated with process control of the SRU but it used to monitor the operation. A key to reliable operation of burner and furnace instrumentation is to adequately purge each instrument nozzle even when the SRU is shut down. The process gas contains elemental sulfur, which will condense and solidify if allowed to enter the cool instrument nozzles. The temperature in the reaction furnace should be monitored. This is a valuable tool to determine if conditions have changed and effected combustion. An optical pyrometer is the best solution. The most critical burner instruments are the flame scanners. If the flame scanners are unreliable, it can result in numerous unwarranted, nuisance shut downs. SRU's require special flame scanners because the H2S flame provides weak ultraviolet radiation. Depending on the burner, some scanners may work and others will not. Reliable scanners operation requires proper installation which includes sighting, grounding, calibration, etc.

6.2.4 Waste Heat Boiler

The level controls associated with the WHB are critical. High pressure boilers should normally utilize more sophisticated three element level control which uses steam production rate and boiler level to set the BFW flow control loop. The three-element control system reacts quicker to changes in load on the boiler and allows the steam drum to be smaller. WHB's also must have low level switches that will initiate an SRU shutdown. Malfunctions in the level control/shutdown system can result in WHB tube failure and major SRU downtime.

6.2.5 Sulfur Condensers

The only process control associated with sulfur condensers is generally simple level control. There are no shutdown switches normally associated with sulfur condensers.

6.2.6 Reheaters

Most reheater types will utilize a slightly different control system. All reheaters will employ a temperature control loop; the differences occur in the stream actually controlled.


96

Direct fired line burners require the most complicated controls and includes automatic shutdowns associated with the firing controls.

6.2.7 Catalytic Reactors

Varying the reheater outlet temperature controls catalytic reactor inlet temperature. There are no other controls associated with the reactors. The installation of thermocouples to monitor bed temperatures, flow distribution, and catalyst activity is also monitored.

6.2.8 Shutdown System

One of the primary improvements that can be made in many SRU's is the installation of a reliable safety shutdown system. The purpose of the SRU shutdown system is to:

Safely shutdown the SRU to protect personnel and/or equipment if critical operating variables move outside the save operating range.

Insure a safe, orderly start-up of the SRU.

Operation personnel tend to do whatever is necessary to keep the SRU on-line even though it can often result in unsafe operating conditions if the system is properly designed, width properly selected and installed equipment, and is properly maintained, and the shutdown system can perform its intended function reliably without nuisance shutdowns. The following items should initiate an SRU Shutdown in all SRU's:

1. One or more operator activated manual switches. 2. High level in amine acid gas knockout drum. 3. Low amine acid gas flow rate. 4. Low combustion air blower discharge pressure. 5. Low main combustion air flow. 6. Loss of main burner flame as detected by two of two flame Scanners. 7. Low waste heat boiler water level.

6.3 Instrumentation for Condensers

Level of boiling feed water in the condenser 1, 2 and 3 is measured by a FEEDBACK Control system. Similarly the pressure of steam exiting from the condenser is kept at 45 psig using a FEEDBACK control system.

CONDENSER 4

Temperature of exiting acid gas stream from the final condenser is controlled by BFW and the temperature of BFW entering the final condenser is controlled by 45 psig steam at temperature of 289 F. Two feedback control system do both the duties.


97

CONDESNER 1-2-3

Level of boiling feed water in the condenser 1, 2 and 3 is measured by a FEEDBACK Control system. Similarly the pressure of steam exiting from the condenser is kept at 45 psig using a FEEDBACK control system.

Chapter 7 MECHANICAL DESIGN

7.1 Introduction

The corrosion problems can be minimized with proper metallurgy. In the SRU, we will discuss the material of construction and lining of refractory for the most important equipments. We will discuss these in the light of sulfide corrosion. Many types of equipment in the SRU highly demand the refractory linings due to many reasons. One of them is the high corrosion of the H2S and Sulfur. Another is to reduce the heat loss to the surroundings. The material of construction and refractory linings for the equipments:

Thermally Reaction Furnace

Waste Heat Boiler

The Claus Reactors

Sulfur Condensers

Sulfur Pits

Thermally Reaction Furnace

The specific features and residence time required for an individual reaction furnace are dependent on several factors including the operating temperature and expected feed impurities.

The burner and furnace are typically constructed of carbon steel. The very severe high temperature, reducing atmosphere process operating conditions in the burner and furnace require protection of the carbon steel.

The carbon steel must be protected from acid corrosion (liquid phase H2O, SO2, SO3) and high temperature sulfide corrosion by H2S. the primary form of protection is insulating the steel from the processing atmosphere with a multi-layer refractory lining. The refractory lining should be designed to maintain the inside metal temperature between 400-600 oF.

Chapter 7 Mechanical Design

99

Experience has shown this temperature range to be adequate to avoid both excessive acid and sulfide corrosion. The normal furnace operating temperature is 1800-2000 oF; however, the hot face material must be capable of withstanding temperature of 2500-3000 degree F which can occur during start-up firing of natural gas.

We feel all new SRU reaction furnace linings should utilize at least 90% alumina hot face material. Brick linings are more durable than cast able or plastic ram materials. The initial installation of brick many be slightly more expensive, but brick normally provides a longer lining life and requires less maintenance. The brick used for the hot face should be mullite-bonded material. The high alumina hot face brick is very dense, has good high temperature strength, but has poor insulating properties.

The second refractory layer should be low iron, class 2600 or 2800 insulating firebrick (IFB). The IFB is a lower density, greater porosity brick than the hot faced high alumina brick, and therefore, has better insulating properties. Insulating castable refractory is also used as the backup layer to the hot-faced brick. If insulating castable is used, it is important to insure the inner surface of the castable is perfectly round. Any out of roundness makes installation of the outer brick layer more difficult and can result in an uneven installation that may have voids between the layers.

If and IFB layer is used, the outer layer (layer next to the shell) is often a 1/8-1/4” layer of ceramic fiber paper. This material has very good insulating properties. Often the ceramic fiber paper provides more insulation that the two layers of brick.

The reaction furnace internal features mentioned above (choke ring, checker, wall, etc.) must be incorporated into the refractory lining design. Waste Heat Boiler

The WHB tube sheet and tube inlets have traditionally been protected using high

alumina ceramic tube ferrules/insert and approximately 4” of 60+% alumina castable refractory on the tubesheet face between the ferrules.

Once installed, the castable is cured into a solid block containing the ferrules. Properly curing the castable is often difficult. It is difficult to maintain the recommended time-temperature profile by firing the SRU main burner and operations is normally being pushed to get the SRU back on-line after a shutdown. Inadequate/improper curing generally results in reduced WHB tubesheet refractory effectiveness and life.

The WHB tube sheet refractory has been an area of high maintenance, particularly is

SRU’s undergoing frequent start-up/shutdown cycles. It is difficult to properly support the


100

castable refractory. Movement in the tubesheet that takes place during cycling and flexing of the tubesheet often cause cracking of ferrules and refractory. Once a crack develops, it allows the hot process gases to reach the tubesheet and/or tube. High temperature sulfide corrosion often occurs, eventually resulting in tube failure. Boiler feedwater (BFW) leakage into the hot refractory lined reaction furnace causes additional damage. Very significant SRU down time is required for retubing and refractory repairs.

For the tube sheets we prefer the austenitic stainless steel, which is more resistant

to the sulfur corrosion. We the newer design thick hex-head ferrules for the WHB tube sheet are a

significant improvement. The hex-head design protects the tube sheet and the tube inlet in one step. The hex-heads fit together to completely cover the timesheet in the tube field area. Ceramic fiber paper is placed between the heads to fill any spaces between heads. The hexhead ferrules can be installed much quicker that the conventional ferrules and castable refractory. The individual ferules can move slightly as the tube sheet moves without causing breakage/cracking. The only castable refractory is outside the tube field near the outer edge of the tube sheet where it is easier to support and is less subject to movement.

The Claus Reactors

Internal thick refractory lining including internal partitions below the top of the

catalyst bed to protect the shell from possible internal sulfur fires.

External insulation for heat conservation and the prevent corrosion of the vessel shell. Catalyst support by refractory covered steel beams that support stainless steel or Alonized steed grating covered with two layers of stainless steel screen. The lower screen should be 4 x 4 mesh, and the upper screen 8 x 8 mesh. Both screens should be tied to the grating using stainless steel wire.

Sulfur Condensers:

Care must be taken in choosing the most appropriate material of construction for the sulfur condensers because of the acid gas is in contact with in the sulfur tubes. The most favorable material of construction for the tube side is the typical austenitic stainless steel which provides a best solution for reducing the corrosion. The shell is usually insulated for reducing the heat loss to the surroundings. Sulfur Pits:

Product sulfur is normally collected in a below grade, concrete pit equipped with steam coils to keep the sulfur molten. The pit does not directly affect the SRU process operation


101

until the SRU must be shut down because of problems with the pit. Some common sulfur pit problems are steam coil leakage, sulfur pumps failure, internal sulfur fires, and even internal explosions. There are a few design features that will significantly improve the reliable operation of the sulfur pit. Construct the pit using sulfate-resistant concrete with limestone-free aggregate. Use piping for the steam coil steam coil steam supply down comers and condensate risers, and any internal components such as ladder rungs that will be alternately covered with liquid sulfur and then exposed to air as the pit level changes.

Chapter 8 HAZOP STUDY OF THE SULFUR RECOVERY UNIT (SRU)

8.1 Introduction

The objective of overall plant safety is to provide safety, health and care to all the manpower working in the plant and to provide maintenance and safety to the plant by reducing risks of accident.

8.2 General Safety Rules

1. Goggles or face shield, gloves will be worn every time a corrosive liquid is sampled or when there is danger of coming in contact with a corrosive liquid.

2. A constant lookout for any condition that would prove hazardous to personnel or equipment should be kept.

3. High-pressure leaks should be reported immediately. 4. Fresh air mask or breathing apparatus should be worn by anyone entering a

hazardous atmosphere. Hazardous environment could include sulfur dioxide, hydrogen sulfide or sulfur vapors.

5. If any chemical spill occurs, clean the area of any source of ignition and wash down the area slowly with water to minimize vaporization.

6. Operators should be aware of location of emergency and safety equipment. 7. Operators should be aware of all the toxic, corrosive and flammable materials

used in the process. 8. Fire fighting equipment should not be tampered with. 9. Access to ladders, escapes, safety showers, eyewash stations and air mask

stations must be kept clear. Waste material and refuses must be kept in proper locations where they will not cause fire.

10. If any safety equipment is not working properly, it should be reported and set right.

11. No smoking is allowed in the plant or its vicinity.

8.3 Building and Process Equipment Safety

Any acceptable design must contain the minimization of building and equipment hazards such as corrosion, fire explosion and hazards caused due to fumes and poisonous materials. Special care should be given to the disposal of waste material. Elimination of process

Chapter 8 HAZOP Study of the SRU

103

leakage and spillage hazards due to corrosion and other factors should be paid extra attention.

8.3.1 Lights

Proper light arrangements should be made, to facilitate the movement and working of personnel and minimizing the risk of tripping over pipes.

8.3.2 Electrical and Mechanical Hazards

Electrical and mechanical hazards should be minimized. Every machine should be equipped with proper safety guards. Poor and faulty wiring, overloaded circuits, and improperly loaded circuits should not be used.

8.3.3 Chemical Hazards

Special care should be given to avoid any exposure to sulfur dioxide and hydrogen sulfide. Most reliable manuals for safety are chemical safety data sheet compiled by manufacturing chemists association. These manuals discuss the safe handling of most hazardous chemicals while also providing drawing, data tables and graphs.

8.3.4 Fire Prevention and Control

The term fire protection applies to that phase of process design which minimizes the fire hazards inherent in the process. Fire control refers to control the production and spreading of fire which has already been started. Minimum safe practices prescribed by nationally recognized fire protection associations, engineering authorities and organizations should be exercised.

8.3.5 Personnel Safety

Special equipment such as safety goggles, gloves, ear muffs etc. should be provided to the personnel so they can enhance their performance in a safe and healthy environment.

8.4 Claus Process

In Claus process we treat acidic and corrosive gases so safety should be kept in mind. Hugh temperature and liquid sulfur are the major concern for safety purposes. Special care

Chapter 8 HAZOP Study of the SRU

104

should be given to avoid any exposure to Sulfur Dioxide, Hydrogen Sulfide and Sulfur. High pressure steam is also an important factor for safety consideration.

8.4.1 Special Hazards and Precautions

1. Sulfur dust explosion occurs with very rapid discharge of flame and pressure waves. When confined to a building, pressure waves can cause a great deal of damage.

2. Sulfur reacts violently with strong oxidizing agents such as nitrates and chlorates. It will undergo chemical change at moderate rates with alkalis.

3. Ground sulfur is usually non-toxic through skin contact, ingestion and inhalation. However, it can irritate the skin and eyes as well as the respiratory system.

4. Sulfur deposited on skin can be washed away with mild soap and water. Eyes that have come into contact with sulfur dust are to be flooded with water for at least 15 minutes.

5. Incipient fires in sulfur storage piles can be frequently smothered by gently shoveling more sulfur, sand or fine earth to exclude all air. For larger fires, water applied as fine mist is the most useful agent. High pressure water sprays disperse the dust into air and should not be used.

6. Steam or noble gases are excellent gases for use in containers that can be closed tightly. Care should be taken that the sulfur dust is not scattered into the air.

7. Gas masks approved for acid gases would not provide adequate protection in a serious sulfur fire.

8. Spills should be handled according to the physical state of sulfur: Molten - Burn danger, containment of spill Solid – fire and explosion danger Flowable – less dangerous, containment of spill.

9. Protective Equipment should be used during the following procedures. - Manufacture or formation of product. - Repair and maintenance of contaminated equipment. - Clean up of leaks and spills. - Any situation resulting in hazardous exposure.

Chapter 9 ECONOMIC ANALYSIS & COST ESTIMATION

9.1 Introduction

A capital investment is required for any industrial process and determination of the

necessary investment is an important part of a Plant Design Project in cost analysis of an

industrial process, Capital investment costs, manufacturing costs and general expenses

including income taxes are taken into consideration.

9.2 Fixed and Working Capital

Fixed capital is the total cost of the plant ready for startup. It includes the cost of:

1. Design and engineering and construction supervision.

2. All items of equipment and their installation.

3. All piping, instrumentation and control system.

4. Building and structure.

5. Auxiliary facilities, such as utilities, land and civil engineering work.

Working capital is the additional investment needed, over and above the fixed

capital, to start the plant up and operate it to the point when income is earned.

It includes the cost of:

1. Start up.

Chapter 9 Economic Analysis & Cost Estimation

106

1. Raw material investment.

2. Finished product investment.

3. Funds to cover outstanding accounts from customers.

Most of the working capital is recovered at the end of the project.

The total investment needed for a project is the sum of the fixed and working capital.

9.3 Total Product Cost

Another equally important part is the estimation of costs for operating the plant

and selling the products.

It is generally divided into following categories:

9.3.1 Manufacturing Cost

It includes

Direct production cost:

Raw materials, operating labor, maintenance, steam, fuel, water etc.

Fixed charges:

Depreciations, taxes etc.

Plant Over Head Costs:

Medical, safety and protection, recreation etc.

9.3.2 General Expenses

It includes

Administrative expenses.


107

Marketing expenses.

R & D.

Financing.

Total capital investment is estimated on the basis of purchase equipment cost. Cost

data given in the “Plant Design and Economics for Chemical Engineers” by Max .S .Peter and

Klaus D. Timberhead, & Chemical Engineering. Vol. 6 by J.M Coulson & J.F Richardson.

The value of Marshall & Swift Equipment cost index are given as

Marshall & Swift cost index value, January 1998 = 1070

Marshall & Swift cost index value, January 2004 = 1218

The present cost purchased of the equipment is calculated as follows:

Present cost = original cost *

timeoriginalatvalueindex

timepresentatvalueindex

....

....

9.4 Equipment Cost

Muffle furnace (F-100)

Mass of austenitic steel used for the shell = 8572 1bm

Mass used for heads = 20% = 0.2 * 8572 = 1714.4 1bm

For man holes and supporting = 15% = 1285.8 1bm


108

Mass used for vessel = 11572.2 lbm = 5.26 ton

Austenitic steel cost = 3700 £ / ton

Refractory cost = 30% = 0.3 * 19462 = 5838.6 £

Total furnace cost = 25300.6 £

Boiler (B-100)

Boiler cost can be given as

Ce = CSN (from volume 6)

Cost constant = C = 60 £

Cost index = n = 0.8

Size parameter = S = 5727.27 kg /hr stream

Purchased equipment cost (1998) = Ce = 60884.87

Cost of boiler = cost (1998)



....

....

= (60884.87)(1218/1070)

= 69306.33 £

Reactor (R-100, R-101, R-102)

Ce = C S “

C = 9300 &

n = 0.40

S = 31.63 cu m


109

Cost of reactor = cost (1998)



....

....

= (37028.07) (1218/1070)

= 42149.7 £

Refractory cost = 30% = 0.3 * 42149.7

= 12644.9 £

Catalyst cost = 50 = 0.5 * 42149.7

= 21074.85 £

Cost of one reactor = 75869.45 £

There are three units so

Reactor cost (R–100, R-101, R-102) = 3* 75869.45

= 227608.35 £

Condenser E-106

Given as

Purchased cost = bare cost * type factor * pressure factor heat transfer area =

125.799 sq m

From fig vol 6

Bare cost = 68* 1000=68000 £

Type factor = 0.8

Pressure factor = 1


110

Purchased cost = 54400 £

Cost of condenser (C-d) = cost (1998)

.....

....



= 54400 (1218 / 1070)

= 61924.5 £

As the heat duty of condenser E-100, E-102, E-104 and E-106 is more, we take 10% more

heat transfer area for the cost calculations, so

Heat transfer area for one condenser = 137 sq m

Bare cost = 70 * 1000 = 70000 £

Type factor = 1.3

Pressure factor = 1

Purchased cost = 91000

Cost of one condenser = cost (1998)

timeoriginalvalueindex


...

....

= 91000 (1218/1070)

= 103586 £

Cost of three condensers = 3* 103586 = 310760.75 £

Total cost = 694900.53 £

For others equipment like pre heaters, knock out drums and other minor, we take factor

20% of the estimated equipments

Total cost = 1.20* 694900.53 = 833880.63 £


111

Total purchased equipment cost = 833880.63 £

Now using ratio factors for estimating capital investment items based on delivered

equipment cost as given in plant design and economics for chemical engineers book.


112

Components % age of purchased equipment Cost (&)

Total purchased cost (E) 100%E

Purchased equipment (installation) 47%E

Instrumentation (installed) 18%E

Piping (installed) 66%E

Electrical (installed) 11%E

Buildings (including services) 18%E

Yard improvement 10%E

Service facilities 70%E

Land 6%E

Total direct plant cost (D) 2885227.02

Engineering & supervision 33%E 275180.61

Construction expenses 41%E 341891.06

Total direct and indirect cost (D+I) 3502298.687

Contractors fee 5% (D+I) 175114.93

Contingency 10% (D+I) 350229.87

Fixed capital investment (F) 4027643.486

Working capital investment 15% F 604146.523

Total capital investment 4631790.01

Total capital investment required = 4631790.01 £

= 495601531 Rs.

REFERENCES

1. R.E. Kirk & D.E. Othmer; “Encyclopedia of Chemical Technology”,

Vol-11

2. Katz, D.L; Vary, J.A; “Hand Book of Natural Gas Engineering”; Mc

Graw Hill Co; Inc., 1959.

3. Private Communication, Robert J. Huston, Chairman Air Permits

Division, Texas Natural Resource Conversion Division, Sulfur Recovery Units,

P.O Box 13087 Austin TX 78711-3087

4. Journal, “Assessment of Recovery Capabilities, 2004” By J.F

Strickland, Michael Quinlan, D. Velasquez, Kellogg Brown & Root Houston,

Texas & Dennis Leppin Howard Meyer GTI Chicago, Illinois

5. U.S. Geological Survey, Mineral Commodity Summaries, February

2000

6. Claus-Plant Combustor Survey, November 2001 By Dan Banks, P.E.

Banks Engineering Inc. Tulsa, Oklahoma

7. “Improving Claus Sulfur Recovery Unit Reliability Through

Engineering Design” By Steve Fenderson, Goar, Allison, & Associates, Inc.

P.O. Box 131390 Tyler, Texas 75713 September 15-18, 1998

8. Perry’s Chemicals Engineers’ Handbook, Vol. 1 & 2, 7th Edition by

Robert H. Perry, Don Green

9. Kern D.Q. “Process Heat Transfer”, McGraw Hill Bool Co., New York,

1950

10. McCabe, W.L & Smith, J.C., “Unit Operation for Chemical

Engineering”, 6th Edition McGraw Hill Book Co., New York

11. Peters, M.S. & Timmerhause, K.D., “Plant Design & Economics or

Chemical Engineers” 2nd Edition McGraw Hill Book Co., New York

12. “Sulfur Recovery.” GPSA Engineering Data Book vol. II, eleventh

edition, 1998. Chapter 22

13. C. Rueter, “CrystaSulf® Process Fills Mid-size Niche for Sulfur

Recovery in Multiple Applications,” Gas TIPS Winter 2002, p.7-12

14. S. Lynn et al, “New Approach to Sulfur Removal Could Reduce

Costs,” Gas TIPS winter 2002, p. 13-20

15. J. Strickland and D. Velazquez, “Assessment of Recovery Capabilities

and Costs of Tail Gas Clean-Up Processes,” GRI Topical Report, February

2000, Unpublished.

16. L. Connock, “Enhanced Sulfur Recovery,” Sulfur, May-June 2003,

No. 286

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