Ismail Project

7/28/2019 Ismail Project

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ABSTRACT

This project is about the development & design of Refinery gas sweetening

system of CPCL-plant 12.

Refinery gas sweetening is the process used to remove the so called Acid Gases

which are hydrogen sulphide and carbon dioxide from the refinery gas streams. These acid gas

removal processes used in the refinery are required either to purify a gas stream for further use in

a process or for environmental reasons.

Though there are various processes meant for the sour gas removal the Chemical

reaction type is selected and adopted on grounds that it can selectively remove H 2S from all the

sour gas components.

For serving this purpose, the newly developed MDEA solution of 35 % mol is

used in a mole ratio of 2.2. The factors that are all considered for the selection of MDEA over

MEA, DEA and DGA are also discussed.

For the entire flow sheet the material and energy balances are developed and the

corresponding equipment are designed.

Plant layout, cost estimation and safety measures are going to be carried out.


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TABLE OF CONTENTS

CHAPTER TITLE PAGE NO.

ACKNOWLEDGEMENT ERROR!BOOKMARK NOTDEFINED.

ABSTRACT I

TABLEOFCONTENTS II

LISTOFTABLES VI

LISTOFFIGURES VII

LISTOFSYMBOLSANDABBREVIATIONS VIII

1 INTRODUCTION

1.1ORGANISATION PROFILE 11.2CRUDE OIL 21.3CRUDE OIL REFINERY 3

1.4REFINERY OPERATIONS 4

2 PROPERTIES

2.1PHYSICAL PROPERTIES OF MDEA 62.2PHYSICAL COMPARISION OF MEA AND MDEA 7

2.3CHEMICAL PROPERTIES OF MDEA 8

3 TYPESOFPROCESSES

3.1VARIOUS PROCESSES FOR THE REFINERY GAS TREATING 93.2REACTION TYPE GAS TREATING 9

3.3PHYSICAL SOLVENT GAS TREATING 10

3.4PHYSICAL/CHEMICAL TYPE 113.5CARBONATE TYPE 12

4 SELECTIONOFPROCESS

4.1SELECTION PROCEDURE FOR THE CHEMICAL REACTION

TYPE FOR GAS SWEETENING 134.2PROCESS PARAMETERS AND DESCRIPTION OF VARIOUS

AMINE SOLVENTS 14

4.3FACTORS THAT ARE CONSIDRED FOR THE SELECTIONOF MDEA OVERMEA,DEA & DGA 19

5 PROCESSFLOWDIAGRAM 20


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6 PROCESSDEVELOPMENTANDDESCRIPTION

6.1ABOUT REFINERY GAS SWEETENING 21

6.2PROCESS DESCRIPTION 22

6.3PROCESS CHEMISTRY 226.4THERMODYNAMIC PRINCIPLES 23

7 MATERIAL BALANCE

7.1FLOW RATE OF COMPONENTS OF SOUR GAS THAT DISTILLEDOUT FROM THE CDU-2 24

7.2MATERIAL BALANCE ACROSS ABSORBER 257.3MATERIAL BALANCE ACROSS STRIPPER 26

8 ENERGYBALANCE

8.1HEAT CAPACITY CONSTANTS FOR THE COMPONENTS OF

SOUR GAS 29

8.2ENERGY BALANCE ACROSS THE ABSORBER 30

8.3ENERGY BALANCE ACROSS THE SOLUTION INTERCHANGER338.4ENERGY BALANCE ACROSS THE COOLER 35

8.5ENERGY BALANCE ACROSS THE STRIPPER 36

8.6CONDENSER DUTY 39

9 HEATEXCHANGERDESIGN

9.1PLANT DATA 419.2PROPERTIES OF THE FLUID 42

9.3CALCULATION OF P&R 439.4FILM COEFFICIENT - TUBE SIDE-LEAN AMINE 459.5FILM COEFFICIENT -SHELL SIDE-COOLING WATER 46

9.6 TUBE WALL RESISTANCE 47

9.7OVERALL HEAT TRANSFER COEFFICIENT 47

9.8HEAT TRANSFER AREA REQUIRED 48

9.9PRESSURE DROP TUBE SIDE 489.10PRESSURE DROP SHELL SIDE 49

9.11DESIGN SUMMARY 50


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10 STRIPPERDESIGN

10.1FINDING THE NUMBER OF TRAYS

(FROM KREMSERS CORRELLATION GRAPH) 51

10.2FINDING THE EQUILIBRIUM DISTRIBUTION RATIO FORINDIVIDUAL COMPONENTS 51

10.3CALCULATING THE IDEAL NUMBER OF TRAYS 52

10.4CALCULATING THE ACTUAL NUMBER OF TRAYS 5210.5KREMSERS CORRELATION GRAPH 5310.6THE INTERNAL DIAMETER OF THE STRIPPING COLUMN 54

10.7CALCULATION OF COLUMN HEIGHT 59

10.8DESIGN SUMMARY 60

11 COSTESTIMATION

11.1INTRODUCTION 6111.2TOTAL CAPITAL INVESTMENT 6111.3CALCULATION OF DIRECT COST FACTOR 62

11.4CALCULATION OF INDIRECT COST FACTOR 63

11.5ESTIMATION OF THE CAPITAL INVESTMENT IN THEAUXILIARY SERVICE,IA 64

11.6ESTIMATION OF THE CAPITAL INVESTMENT

AS WORKING CAPITAL,IW 65

11.7ESTIMATION OF TOTAL CAPITAL INVESTMENT 65

12 PLANTLOCATIONANDLAYOUT

12.1PLANTLAYOUT 66

12.2FACTORSINPLANNINGLAY-OUT 66

12.3METHODSOFLAY-OUTPLANNING 66

12.4 SCALE MODELS 66

12.5 EQUIPMENT LAYOUT 68

12.6 PLANT LOCATION 69

13 HAZARDSANDSAFETY

13.1SAFETY 7213.2HAZARDS IN REFINERIES 7213.3DESIGN SAFETY 72

13.4LAYOUT 73

13.5OPERATING PROCEDURES 7313.6STATIC EQUIPMENT 73

13.7FIRE 73

13.8HEALTH HAZARD 74


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13.9IMPORTANT SAFETY RULES TO PREVENT HIGH

PRESSURE HAZARDS 74

13.10ENVIRONMENTAL CONTROL 75

14 CONCLUSION 76

REERENCES 77


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LIST OF TABLES

Table 1. 1Composition of Crude Oil 2Table 1. 2 Composition of Hydrocarbons 3

Table 1. 3 Physical comparision of MEA and MDEA 7

Table 7. 1 Flow rate of components of sour gas thats distilled out from the CDU.2 24

Table 7. 2 The total liquid and vapour flow rates 56

Table 8. 1Heat capacity constants for the components of sour gas 29

Table 8. 2 Enthalpy of sour gas entering at 314k 30

Table 8. 3 Enthalpy of sweetened gas leaving absorber at 311k 31

Table 8. 4 Enthalpy of strip out gases at 380 k 38

Table 8. 5 Enthalpy of Acid gas to SRU at 313K: 39

Table 11. 1 Estimation of fixed capital investment 62

Table 11. 2 Calculation of direct cost factor 63

Table 11. 3 Calculation of indirect cost factor 63

Table 11. 4 Estimation of the capital investment in the auxiliary service 64


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LIST OF FIGURES

Figure 6. 1 Process flow diagram 20

Figure 7. 1 Material balance across absorber 25

Figure 7. 2 Material balance across the stripper 26

Figure 8. 1 Energy balance across the absorber 30

Figure 8. 2 Energy balance across the solution interchanger 33

Figure 8. 3 Energy balance across the cooler 35

Figure 8. 4 Energy balance across the stripper 36

Figure 8. 5 Condenser duty 39

Figure 9. 1 Hot amine cooler 42

Figure 9. 2 Graph for finding the correction factor 43

Figure 10. 1Kremsers correlation graph 53

Figure 10. 2 Graph for finding the Brown and Souders flood constant 58

Figure 12. 1 Plant layout 67


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LIST OF SYMBOLS AND ABBREVIATIONS

LIST OF SYMBOLS

Symbols Name Units

T Difference In Temperature C

TLMTD Logarithmic Mean Temperature Differenc C

Tt True Mean Temperature Difference C

Ft Temperature Correction Factor

Degree

Di Inlet Diameter m

do Inlet Diameter m

L Length of Tube m

PT Square Pitch m

De Equivalent Diameter m

Db Bundle Diameter m

Ds Shell Diameter m

B Baffle Spacing m

ut Tube Velocity m/s

us Shell Velocity m/s

Viscosity kg/m s

Cp Specific heat capacity kJ /kgoC

T Temperature C or KTo Temperature - Out C or KTi TemperatureIn C or K


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QH Heat Transfer dutyHot fluid kcal/h

QC Heat Transfer dutyCold fluid kcal/h

U Overall Heat Transfer Coefficient W / m2

K

H Heat Transfer Coefficient W / m2

K

D Diameter m

L Length m

A Area m2

m Mass flow rate kg / h

k Thermal Conductivity W / moC

Density kg / m3

NTU Number of Transfer Unit m

LP Length of the plate m

LW Width of the plate m

DP Diameter of the plate m

b Thickness of the plate m

AP Area of single plate m2

As Effective area of Shell m2

Xm Thickness of Material m

Nt

Total number of tubes

NP Number of passes

Re Reynolds Number

Pr Prandtl Number

Nu NussultNumber


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LIST OF ABBREVIATIONS

CPCL - Chennai Petroleum Corporation Limited

MRL - Madras Refinery Limited

GoI - Government of India

AMOCO - American Oil Company

NIOC - National Iranian Oil Company

IOC - Indian Oil Corporation Limited

EIL - Engineers India Limited

MMTPA - Million Metric Tonnes Per Annum

WHO - World Health Organization

FAO - Food and Agriculture Organization

UNISEF - United Nations Childrens Fund

LSRG - Light Straight Run Gasoline

NMP - Normal Methyl Pyrrolidone

MEA - Mono Ethanol Amine

MDEAMethyl di ethanol Amine

DEADi ethanol Amine

DGADi glycol Amine

SRC - Solvent Recovery Column

IBP - Initial Boiling Point

TBP - Total Boiling Point

MP - Medium Pressure

LP - Light Pressure

HP - High Pressure

SHE - Spiral Heat Exchanger

LEL - Lower Explosive Limit


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CHAPTER 1

INTRODUCTION

1.1 Organisation profileChennai Petroleum Corporation Limited (CPCL), formerly known as Madras Refineries

Limited (MRL) was formed as a joint venture in 1965 between the Government of India (GOI),

AMOCO and National Iranian Oil Company (NIOC) having a shareholding in the ratio

74%:13%:13% respectively. Originally, CPCL Refinery was set up with an installed capacity of

2.5 Million Tonnes Per Annum (MMTPA).

CPCL has two refineries with a combined refining capacity of 11.5 Million Tonnes Per

Annum (MMTPA). The Manali Refinery has a capacity of 10.5 MMTPA. CPCL's second

refinery is located at Cauvery Basin at Nagapattinam. The Manali Refinery located at Chennai is

one of the most complex and integrated refineries with three crude distillation units, Diesel

Hydro De-sulphurisation unit, Fluid Catalytic Cracking unit, Furfural Extraction unit, Lube

Hydrofinishing unit, NMP Extraction unit, Hydro-Cracker unit, Propylene unit and

Petrochemical Feedstock unit.

The main products of the company are LPG, Motor Spirit, Superior Kerosene, Aviation

Turbine Fuel, High Speed Diesel, Naphtha, Bitumen, Lube Base Stocks, Paraffin Wax, Fuel Oil,

Hexane and Petrochemical feed stocks. CPCL, ever since its inception, has been methodically

planning and implementing several environmental conservation measures. A dedicated

Environment Management Team functions exclusively to plan, implement, operate and monitor

all environment-related activities.

CPCL Manali Refinery has obtained ISO 9001, ISO 14001 and OHSAS-18001

certifications.


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1.2 Crude oilCrude Oil is a naturally occurring flammable liquid consisting of a complex mixture

ofhydrocarbons of various molecular weights and other liquid organic compounds, that are

found in geologic formationsbeneath the Earth's surface. Petroleum is recovered mostly

through oil drilling. The use of fossil fuels such as petroleum can have a negative impact on

Earth's biosphere, releasing pollutants and greenhouse gases into the air and damaging

ecosystems through events such as oil spills. Concern over the depletion of the earth's finite

reserves of oil, and the effect this would have on a society dependent on it, is a field known

as peak oil.Composition of Crude Oil

In its strictest sense, petroleum includes only crude oil, but in common usage it includes

all liquid, gaseous, and solid hydrocarbons. The exact molecular composition varies widely from

formation to formation but the proportion ofchemical elements vary over fairly narrow limits.

Elements Composition (%)

Carbon 83 to 87

Hydrogen 10 to 14

Nitrogen 0.1 to 2

Sulphur 0.05 to 6

Oxygen 0.05 to 1.5

Metals Traces

Table 1- 1Composition of Crude Oil
http://en.wikipedia.org/wiki/Flammable_liquidhttp://en.wikipedia.org/wiki/Hydrocarbonhttp://en.wikipedia.org/wiki/Organic_compoundhttp://en.wikipedia.org/wiki/Formation_(stratigraphy)http://en.wikipedia.org/wiki/Earthhttp://en.wikipedia.org/wiki/Oil_drillinghttp://en.wikipedia.org/wiki/Oil_spillhttp://en.wikipedia.org/wiki/Oil_depletionhttp://en.wikipedia.org/wiki/Non-renewable_resourcehttp://en.wikipedia.org/wiki/Non-renewable_resourcehttp://en.wikipedia.org/wiki/Peak_oilhttp://en.wikipedia.org/wiki/Hydrocarbonshttp://en.wikipedia.org/wiki/Chemical_elementhttp://en.wikipedia.org/wiki/Chemical_elementhttp://en.wikipedia.org/wiki/Hydrocarbonshttp://en.wikipedia.org/wiki/Peak_oilhttp://en.wikipedia.org/wiki/Non-renewable_resourcehttp://en.wikipedia.org/wiki/Non-renewable_resourcehttp://en.wikipedia.org/wiki/Oil_depletionhttp://en.wikipedia.org/wiki/Oil_spillhttp://en.wikipedia.org/wiki/Oil_drillinghttp://en.wikipedia.org/wiki/Earthhttp://en.wikipedia.org/wiki/Formation_(stratigraphy)http://en.wikipedia.org/wiki/Organic_compoundhttp://en.wikipedia.org/wiki/Hydrocarbonhttp://en.wikipedia.org/wiki/Flammable_liquid


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Four different types of hydrocarbon molecules appear in crude oil. They are Paraffins,

Napthalenes, Aromatics and Asphaltics.

Hydrocarbon Average

Composition (%)

Paraffins 30

Napthalenes 49

Aromatics 15

Asphaltics 6

Table 1- 2 Composition of Hydrocarbons

Crude oil varies greatly in appearance depending on its composition. It is usually black or

dark brown. In the reservoir it is usually found in association with natural gas, which being

lighter forms a gas cap over the petroleum, and saline waterwhich, being heavier than most

forms of crude oil, generally sinks beneath it. Crude oil may also be found in semi-solid form

mixed with sand and water, as in the Athabasca oil sands in Canada, where it is usually referred

to as crude bitumen.

1.3 Crude oil refineryAn oil refinery or is an industrial process plant where crude oil is processed and refined

into more useful products such as petroleum naphtha, gasoline, diesel fuel, asphalt

base, kerosene, and liquefied petroleum gas. Oil refineries are typically largeindustrial complexes with extensive piping running throughout, carrying streams of fluids

between large chemical processing units. The process is very complex and involves both

chemical reactions and physical separations.
http://en.wikipedia.org/wiki/Saline_waterhttp://en.wikipedia.org/wiki/Athabasca_oil_sandshttp://en.wikipedia.org/wiki/Bitumenhttp://en.wikipedia.org/wiki/Bitumenhttp://en.wikipedia.org/wiki/Athabasca_oil_sandshttp://en.wikipedia.org/wiki/Saline_water


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1.4 Refinery operations

Crude Oil DistillationCrude oil distillation is used to separate the hydrocarbons in crude oil into fractions based

on their boiling points. The separation is done in a large tower that is operated at atmospheric

pressure. The tower contains a number of trays where hydrocarbon gases and liquids

interact. The liquids flow down the tower and the gases up. The lighter materials such as butane

and naphtha are removed in the upper section of the tower and the heavier materials such as

distillate and residual fuel oil are withdrawn from the lower section.

Catalytic ReformingCatalytic reforming is a chemical process used to convert petroleum refinery naphthas,

typically having low octane ratings, into high-octane liquid products called reformates which are

components of high-octane gasoline. Basically, the process re-arranges or re-structures

the hydrocarbon molecules in the naphtha feedstocks as well as breaking some of the molecules

into smaller molecules.

Catalytic CrackingFluid catalytic cracking (FCC) is the most important conversion process used in

petroleum refineries. It is widely used to convert the high-boiling, high-molecular weight

hydrocarbon fractions of petroleum crude oils to more valuable gasoline, olefinic gases, and

other products. Cracking of petroleum hydrocarbons was originally done by thermal cracking,

which has been almost completely replaced by catalytic cracking because it produces more

gasoline with a higher octane rating. It also produces byproduct gases that are more olefinic, and

hence more valuable, than those produced by thermal cracking.

Alkylation and IsomerisationIn the alkylation process, isobutane is reacted with either isobutylene or propylene to

form complex paraffin isomers. The reactions take place in the presence of hydrofluoric or

sulfuric acid catalysts. By combing these molecules the octane level of the paraffin isomer or


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alkylate is increased to around 93-96 octane. Refiners use this process to improve the octane

level of the gasoline pool. Light naphtha can have its octane number improved by the use of

an isomerization process to convert normal paraffins into their isomers. This results in an

increase in octane number as evidenced by increase in normal pentane to iso-pentane. The

process uses a platinum catalyst. Like alkylation, this process improves the octane quality of the

gasoline pool.

HydrotreatingHydrotreating is a process where a petroleum fraction is reacted with hydrogen for the

purpose of removing impurities. The process is usually used to remove sulfur. Hydrotreating

processes use hydrogen from the catalytic reformer or a hydrogen plant.

Product BlendingProduct blending is where the different petroleum fractions are combined together to

make the final product. The fractions are mixed so they meet the specifications discussed

earlier. Each product has a specific recipe that calls for the proper mix of petroleum

fractions. For example, in order to make gasoline, the refiner would mix naphtha, reformate,

catalytic gasoline, alkylate and butane so that the mixture had the required octane number, vapor

pressure, sulfur level and aromatics content. The process requires knowing these values for all

of the components going into the blend.


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CHAPTER 2

PROPERTIES

2.1 Physical properties of MDEA

Structure CH3 N (C2H4OH) 2

Molecular weight 119 g/mol

Strength of solution used 35%mol

Boiling point at 760 mm hg - 247.3c

at 50 mm hg - 163.5 c

at 10 mm hg- 128.6c

Refractive index at 20C -1.4694

Heat of combustion at 25c -12,200 btu/lb

Flash point Pensky-martens closed cup (ASTM D93), 138 C

Freezing point -21C

Vapor pressure at 20c, mm hg


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2.2 Physical comparision of MEA and MDEAPreviously in the refinery the MEA solvent had been used since the inception of this

plant. Later on due to the following advantages that MDEA is possessing over MEA the MDEA

has been adopted. The following table gives the physical comparision of MEA over MDEA.

S.No Monoethanolamine (MEA) Methyldiethanolamine (MDEA)

1. Low solvent cost compared to other solvents.

Higher solvent cost relative to

Monoethanolamine, Diethanolamine

and Diglycolamine agent.

2.

High solvent vapor pressure which results in

higher solvent losses than the other

alkanolamines.

Low vapor pressure which results in

potentially lower solvent losses.

3.Higher corrosion potential than other

alkanolamines.

Methyldiethanolamine is less

corrosive.

4.Partial removal of COS and CS2, which

requires a reclaimer.

Non-reclaimable by conventional

reclaiming techniques, and Minimal

COS, CS2 removal.

5.High energy requirements due to the high heat

of reaction.Efficient energy utilization.

6.Nonselective removal in a mixed acid gas

system.

Selectivity of H2 S over CO2 in mixed

gas system.

7.

Formation of irreversible degradation products

with CO2, COS and CS2 which requires

continuous reclaiming.

High resistance to degradation.

8.Molecular weight of Monoethanolamine is

61.1 g/mol

Molecular weight of

Methyldiethanolamine is 119.21 g/mol

10.High reactivity due to its primary amine

character.Lower comparative reactivity.

11.Gas is treated at low pressures, and maximum

removal of H2S and CO2 is required.

Complete H2S removal while only a

portion of CO2 is removed.

12.

More efficient solvents particularly for the

treatment of high pressure natural gas are

rapidly replacing MEA.

The most significant application of

MDEA solvent was in tail gas treating

units.

Table 1- 3 Physical comparision of MEA and MDEA


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2.3 Chemical properties of MDEAReaction of MDEA with H2S

2RNH2 + H2S (RNH3)2S

(RNH3)2S + H2S 2 RNH3HS


The above equations mention the chemistry of the reaction between MDEA and H2S.It can be seen that stoichiometrically 1 mol of MDEA can absorbs 1 mol of H2S from the sour

gas.

Selective removal of H2S over CO2 by MDEA

The biggest advantage that MDEA possess over other amine solvents such as DGA,

MEA, DEA is that it can selectively remove the H2S from the sour gas over the presence of CO2.

Non formation of corrosive heat stable compounds in the presence of COS,

CS2

The MDEA solution doesnt yield a corrosive nature unlike the formation of heat stable

compounds by DGA and MEA in the presence of COS, CS 2. This gives an advantage not to operate a

reclaimer to regenerate the amine solution thereby equipment cost that would be spent is prevented.


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CHAPTER 3

TYPES OF PROCESSES

3.1 Various processes for the refinery gas treatingThe selected popular processes for the refinery gas treating are grouped as follows:

1. Reaction type: MEA, DEA, MDEA, DGA, Stretford

2. Physical solvent type: Fluor solvents (propylene carbonate)

3. Physical chemical type: Sulfinol

4. Carbonate type: Potassium carbonate

3.2 Reaction type gas treatingThe amines (MEA, DEA, MDEA, DGA) are the most popular treating solutions. At one

time MEA, and later MEA and DEA, dominated the market. Amines in general should be

considered for:

Low acid gas partial pressures (product of system pressure and concentration of acidgases-H2S and CO2 in the feed) of roughly 50 psi and below.

Low acid gas concentrations in the gas product of roughly 4-8 ppm. Heavier hydrocarbons present ( provide better filtration for DGA) For CO removal with no H2S, a CO2 partial pressure of 10-15 psi.


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Application of various amine solvents and their respective cases:

Among the amines,

1. MEA is preferred for low contactor pressures and stringent acid gas specificationssuch as H2S well below 0.25grains/ l00scf, and C02 as low as 100 ppmv. MEA is

degraded by COS and CS2 with the reactions only partially reversible with a

reclaimer.

2. DEA is preferred when system pressure is above 500psi. The 0.25 grains/ l00scfis more difficult to produce with DEA. COS and CS 2 have few detrimental effects

on DEA. This, and high solution loadings, provide advantages over MEA. DEA

will typically be used for refinery and manufactured gas streams that have COS

and CS2.

3. MDEA is preferred for selective H2S removal and lack of degradation from COSand CS2.

4. DGA is preferred for cold climates and high (50-70 wt %) solution strength foreconomy. By comparison, solution strength for MEA is 15-25 wt%, and for DEA,

25-35 wt %. Provide good filtration for DGA because it has a greater affinity for

heavy hydrocarbons than other amines. The feed gas must have at least 1% acid

gas for DGA to provide savings over MEA.

3.3 Physical solvent gas treating

The physical solvent types of gas treatment are generally preferred when acid gases in the

feed are above 50- 60psi. This indicates a combination of high pressure and high acid gas

concentration.


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Heavy hydrocarbons in the feed discourage physical solvents, but not COS and CS 2,

which do not degrade the solvents.

Usually, physical solvents can remove COS, CS2, and mercaptans. Physical solvents are

economical because regeneration occurs by flashing or stripping which require little energy.

Selexol, licensed by the Norton Company, uses the dimethyl ether of polyethyleneglycol. A Selexol plant can be designed to provide some selectivity for H2S.

For example, the plant can be designed to provide pipeline quality gas while slipping

85%of the CO.

The Fluor solvent, propylene carbonate, is used primarily for removal of CO, fromhigh pressure gas streams. The CO, off gas stream was used for enhanced oil

recovery.

3.4 Physical/chemical typeThe Sulfinol process from Shell Development Company is a good example of the

physical/chemical type of process. It blends a physical solvent and an amine to obtain the

advantages of both. The physical solvent is Sulfolane (tetra hydro thiophene dioxide) and the

amine is usually DIPA (di isopropanol amine). The flow scheme is the same as for an amine

plant.

Advantages of the Sulfinol Process

The physical solvent Sulfolane like other physical solvents, has higher capacityfor acid gas at higher acid gas partial pressures. At these higher partial pressures

the Sulfinol B process has lower circulation rates (higher solution loading) and

better economy than MEA. Sulfinol is demonstratively advantageous against


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MEA when the H2S/CO2 ratio is greater than 1 : 1 while at high acid gas partial

pressures.

In the solution, the amine DIPA is meanwhile able to achieve pipeline quality gas. COS, CS2 and mercaptans are removed. CO2 slightly degrades DIPA, but

reclaiming is easy.

Low corrosion/carbon steel. Low foaming. Low vapor losses.

3.5 Carbonate typeHot (230-240F) potassium carbonate treating was patented in Germany in 1904 and

perfected into modern commercial requirements by the U.S. Bureau of Mines.

The U.S. Bureau of Mines was working on Fischer-Tropsch synthesis gas at the time.

Potassium carbonate treating requires high partial pressures of CO2. It therefore cannot

successfully treat gas containing only H2S.


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CHAPTER 4

SELECTION OF PROCESS

4.1 Selection procedure for the Chemical Reaction type for gas sweeteningOf the gas treating methods mentioned above, the Chemical reaction type has been

selected on basis of the following reasons:

Potassium carbonate treating requires high partial pressures of CO2. It thereforecannot successfully treat gas containing only H2S.

The physical/ chemical type, the SULFINOL process possess high pricedchemicals and process royalty and the licensee requires many engineering

services.

The physical solvent types of gas treatment are generally preferred when acidgases in the feed are above 50- 60 psi.

Usually, physical solvents can remove COS, CS2, and mercaptans. But the gas

that has to be treated has no presence of mercaptans and else.

Thus the amine solvents- the chemical reaction type can selectively remove H2Sand do not show any degradation by COS, CS2 and so it is selected for the sour

gas sweeteing.


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4.2 Process parameters and description of various amine solvents

Mono ethanol amineThis is the most common acid gas absorption process. Normally 1520 wt% MEA in

water is circulated down through a trayed absorber to provide intimate contact with the sour gas.

The rich solution is routed to a steam stripping column where it is heated to about 250F

at 10 psig to strip out the acid gases. The lean MEA solution is then returned to the absorber.

MEA is the most basic (and thus reactive) of the ethanol amines. MEA will completely

sweeten sour gases removing nearly all acid gases if desired. The process is well proven in

refinery operations.

Like all of the amine solvents used for acid gas removal MEA depends upon its amino

nitrogen group to react with the acidic CO2 and H2S in performing its absorption. The particular

amines are selected with a hydroxyl group which increases their molecular weight and lowers

their vapor pressures yielding minimum solvent losses to the gas stream. MEA is considered a

chemically stable compound. If there are no other chemicals present it will not suffer degradation

or decomposition at temperatures up to its normal boiling point.

The process reactions are given below:

HOCH2CH2NH2 = RNH2 = MEA

Low temp

2 RNH2 + H2S (RNH3)2SHigh temp

Low temp

(RNH3)2S + H2S 2RNH3HS

High temp


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Some of the degradation products formed in these systems are highly corrosive. They are

usually removed by filtration or reclaimer operations.

Filtration will remove corrosive by products such as iron sulfide. Reclaiming is designed

to remove heat stable salts formed by the irreversible reaction of MEA with COS, CS 2 (carbonyl

sulfide and carbon disulfide).

The reclaimer operates on a side stream of 13% of the total MEA circulation. It is

operated as a stream stripping kettle to boil water and MEA overhead while retaining the higher

boiling point heat stable salts.

When the kettle liquids become saturated at a constant boiling point with the degradation

products it is shut in and dumped to the drain. Union Carbide has developed a well proven

corrosion inhibitor system for MEA that allows solution strengths of 2832%. The inhibitor

requires payment of a royalty. The inhibitor chemicals are both expensive and hazardous for

personnel to handle. The system does reduce corrosion problems to nearly zero and allows much

higher system capacity for the same size equipment.

In addition to the chemical degradation mentioned above, MEA oxidizes when exposed

to air. Storage and surge tanks must be provided with inert blanket gases such as N 2 or sweet

natural gas to avoid this degradation.

Amine systems foam rather easily resulting in excessive amine carryover in the contactor.

Foaming can be caused by solids such as carbon or iron sulfide; condensed hydrocarbon liquids

from the gas stream; degradation products; almost any foreign material introduced to the system

such as valve grease, excess corrosion inhibitor, etc. Some of these items such as iron sulfide or

carbon particles are removed by cartridge filters. Hydrocarbon liquids are usually removed by

the use of a carbon bed filter on a lean amine side stream (about 10% of total flow). Corrosion

byproducts are removed by reclaiming as noted above.


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High skin temperatures on the reboiler / reclaimer tubes promote amine degradation.

Steam or hot oil used for the reboiler should be limited to a maximum of 285F (140

C) to avoid

excessive temperatures. The reclaimer should not see hot oil or steam above 415F (213C).

MEA is nonselective in absorbing acid gases. It will absorb H2S faster than CO2 but the

difference is not significant enough to allow its use to separate them. With the lowest molecular

weight of the common amines, it has a greater carrying capacity for acid gases on a unit weight

or volume basis. This generally means less pure amine circulation to remove a given amount of

acid gases. Because the solvent is in solution with water, the gas with which it comes in intimate

contact will leave the cotactor at its water saturation point. If dehydration is necessary it must be

done after the MEA system.

Diethanol amineDEA does not degrade when contacted with CS2, COS, and mercaptans as does MEA.

However for product gas streams which must meet lower than 1 grain per scf of H2S, MEA must

be used. Because of this, DEA has been developed as a preferred solvent when these chemicals

are present in the stream to be treated.

The reaction with acid gas for any of the amines is a mole to mole reaction. The

molecular weight of DEA is 1.7 times that of MEA. Even after correcting for density it requires

1.6 lb of DEA to react with the same amount of acid gas as 1 lb of MEA.

DEA is a weaker base (less reactive) than MEA. This has allowed DEA to be circulated

at about twice the solution strength of MEA without corrosion problems. DEA systems are

commonly operated at strengths up to 30 wt% in water and it is not unusual to see them as high

as 35 wt%. This results in the DEA solution circulation rate usually being a little less than MEA

for the same system design parameters.


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The process reactions are shown below.

HOCH2CH2NHCH2CH2OH = R2NH = DEA

Low temp

2R2NH + H2S (R2NH2)2S

High temp

Low temp

(R2NH2)2S + H2S 2R2NH2HS

High temp

Because the system has much fewer corrosion problems and removes acid gases to

nearly pipeline specifications it has been installed as the predominant system in recent years.

Diglycol amineThis process has been developed by the Fluor Company. It originally began as a

combination of 15% MEA, 80% tri ethylene glycol, 5% water. The system would both sweeten

and dehydrate (to the same level as 95% TEG) the gas in a single step.

The high vapor release during regeneration (both water vapors and acid gases) causes

severe erosion/corrosion problems in the amine/amine exchanger and in the regeneration column.

This system has generally been abandoned.

The present system uses 2- (2-amono ethoxy) ethanol at a recommended solution strength

of 60 wt% in water. DGA has almost the same molecular weight as DEA and reacts mole for

mole with acid gases. DGA seams to tie up acid gases more effectively so that the higher

concentration of acid gas per gallon of solution does not cause corrosion problems as

experienced with the usual amine systems.


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The system reactions are given below.

HOCH2CH2OCH2CH2NH2 = RNH2 = DGA

Low temp


High temp

Low temp

(RNH3)2S + H2S 2RNH3HS

High temp

DGA does react with COS and mercaptans similarly to MEA but forms N, N1, bis

(hydroxy, ethoxy ethyl) urea, BHEEU. BHEEU can only be detected using an infrared testrather than chromatography. Normal operating levels of 24% BHEEU are carried in the DGA

without corrosion problems. BHEEU is removed by the use of a reclaimer identical to that for an

MEA system but operated at 385F (196

C).

There has been a concern that DGA might be a good solvent for unsaturated

hydrocarbons. A survey of the DGA users indicates that many of the systems are operated on gas

containing concentrations of C5+ above 2% without any indication of hydrocarbon loading of

the system.

Those systems near their hydrocarbon dew point are usually installed with a flash tank on

the rich amine from the absorber. The flash tank is operated at a reduced pressure just high

enough to get into the plant fuel gas system. It reduces the vapor load on the regenerator column.

(A similar system is recommended on MEA systems operating near the hydrocarbon dew point.)

DGA allows H2S removal to less than 0.25 grain per 100 scf and removes CO2 to levels

of about 200 ppm using normal absorber design parameters.


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4.3 Factors that are considred for the selection of MDEA over MEA, DEA& DGA

1. The overall reaction of MEA gives us one mole of H2S to one mole of MEA. But thisis seldom experienced in actual contactor due to the limitations in ideal contacting. A

ratio of 3 moles of MEA to 1 mole of acid gas is used.

2. The usage of MEA has receded in the recent years due to its corrosive nature byforming heat stable compounds in the presence of COS & CS2. This requires the use

of a reclaimer to regenerate back the amine solution. But this leads to the loss of

amine and also adds up to the equipment cost.

3. MEA may degrade at higher temperatures around 160c.

4. DEA does not degrade but has lesser absorbing capacity as it is a weaker base.

5. In the case of DGA the high vapor release during regeneration (both water vapors andacid gases) causes severe erosion / corrosion problems in the amine exchanger and in

the regeneration column. So this system has generally been abandoned.


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CHAPTER 5

PROCESS FLOW DIAGRAM

Figure 6- 1 Process flow diagram


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CHAPTER 6

PROCESS DEVELOPMENT AND DESCRIPTION

6.1 About Refinery gas sweeteningRefinery gas sweetening is the process used to remove the so called Acid Gases

which are hydrogen sulphide and carbon dioxide from the refinery gas streams. These acid gas

removal processes used in the refinery are required either to purify a gas stream for further use in

a process or for environmental reasons.

Clean air legislation now being practiced through most industrial countries require

the removal of these acid gases to very low concentrations in all atmosphere to form very diluted

sulphuric acid and carbon dioxide to form carbonic acid both of which are considered injurious

to personal health. These compounds also cause excessive corrosion to metals and metallic

objects.

The removal of sour or acid gas components such as hydrogen sulphide, carbon

monoxide, carbonyl sulphide and mercaptans from gas and liquid hydrocarbon streams is a

process requirement in many parts of the hydrocarbon processing industry.

Typical industrial gas treating application include the production of natural gas,

refinery fuel gas, high purity hydrogen and carbon monoxide, and purification of ammonia,

synthesis gas, landfill gas, coke oven gas or sulphur recovery unit tail gas. For over fifty years,

the alkanolamines have been utilized for acid gas removal in the natural gas and petroleum

processing industries.


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6.2 Process descriptionSour gas (rich in H2S) enters the bottom of the trayed absorber (or Contactor).

Lean Amine is introduced at the top tray of the absorber section to move down the column

operating at 45 MPa g.

Contact between the gas and amine liquid on the trays results in the H2S in the gas

being absorbed into the amine. The sweet gas is water washed to remove any entrained amine

before leaving the top of the contactor. Rich amine leaves the bottom of the contactor at 340.51

k to enter a surge drum. If the contactor pressure is high enough a flash stream of H2S can be

routed from the drum to a trayed stripper.

The liquid is preheated to 459.01 k before entering a stripping column on the top

stripping tray that is operating at 125 kPa a. This stripper is reboiled with 10.5 bar saturated

steam.

The H2S is stripped off and leaves the reflux drum to the sulphur production plant

12 at 313 k. The lean amine leaves the stripper bottom at 389 k and is cooled. The cooled

stream is routed to the contactor.

6.3 Process chemistry2RNH2 + H2S (RNH3)2S

(RNH3)2S + H2S 2 RNH3HS


Though stoichiometrically, MDEA of 1 mol absorbs 1 mol of H2S, for practical cases the

mole ratio is set at 2.2 mol MDEA/ mol H2S.


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6.4 Thermodynamic principles

The process reaction for the absorption of H2S on MDEA is as follows:

CH3 N (C2H4OH) 2= RNH2 = MDEA

Low temp

2 RNH2 + H2S (RNH3)2SHigh temp

Low temp

(RNH3)2S + H2S 2RNH3HSHigh temp

The underlying thermodynamic principle for any reaction at equilibrium is Le Chatliers

principle.

A simple study of the above reaction would indicate that a high pressure and low temperature is

expected for thermodynamic feasibility of absorption of the acid gases.

Thus a high pressure of 45 MPa g is maintained in the absorber and the inlet sour gases are fed at

a low temperature of 314 k.

Similarly, due to the inverse nature of stripping process, a low pressure of the order of 150 kPa is

maintained in the stripper. Here steam at 10.5 bar is provided to strip the acid gases from the

down coming amine solution. A reboiler is used to maintain the temperature profile along the

column. This is limited by the boiling point of the amine solution or the point to which it is not

degradable.


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CHAPTER 7

MATERIAL BALANCE

7.1 Flow rate of components of sour gas that distilled out from the CDU-2

Table 7- 1 Flow rate of components of sour gas thats distilled out from the CDU-2

Component of

sour gaskmol/h

H2S 27.21

H2 36.32

C1 59

C2 35.7

C3 46.6

iC4 18.9

n-C4 20.4

n-C5 1.3

C6 1.6

TOTAL 247.03


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7.2 Material balance across absorber

Figure 7- 1 Material balance across absorber

Total sour gas into the absorber = 247.03 kmol/hAs a mole ratio of 2.2 kmol MDEA/ kmol H2S is set and a solution of strength 35 % mol has

been used,

Total Lean Amine circulation rate forthe absorption of 27.21 kmol H2S = 170.97 kmol/h

Rich Amine rate which has absorbed27.21 kmol H2S = 170.97 +27.21

= 198.18 kmol/h

Sweetened gas flow rate = 219.82 kmol/h


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Overall material balance:

Input = Output

Sour gas + Lean Amine = Sweetened gas + Rich Amine

247.03 + 170.97 = 219.82 + 198.18 kmol/h

7.3 Material balance across stripper

Figure 7- 2 Material balance across the stripper

Rich Ammine entering the stripper = 198.18 kmol/h

500 ppm loading of H2S in Lean Amine = 500 * 10-6

*170.97

= 0.085 kmol/h

Stripped gas contains H2S & water thats condensed and returned as Reflux.

In the accumulator , the H2S leaves at 40C with moisture in it.


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H2S in the stripped gas = 27.21- 0.085= 27.125 kmol/h

Reflux flow rate = 27.125 kmol/h (Since reflux ratio =1)

H2O (entrapped moisture) in the H2S leaving the accumulator:This can be found by making use of partial & vapour pressure data.

This gives the makeup H2O rate.

Considering the contact of Reflux water thats condensed & H2S in the accumulator.

As H2O is pure the vapour pressure @ 40C = Partial pressure of H2O in H2S

leaving accumulator

Now we can obtain the mole fraction of H2O in H2S, using Raoults law.

Pa = ya P ;

where

Pa= partial pressure of the component a in vapour.

P = Total pressure of the gas stream =150-25 = 125 kPa.

Assuming 25 kPa pressure drop in condenser.

ya= 7.375/125 = 0.059

Using the mole fraction obtaining the specific moisture content.

Specific moisture content = M.F. of H2O / M.F. of H2S

= 0.0626 kmol H2O/ kmol dry gas


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The amount of H2O leaving with H2S from the Accumulator

= 0.0626 * 27.125 kmol/h

= 1.70 kmol/h

This is the amount that lost from the reflux water. Thats the make up forreflux.

So a total of 28.90 kmol/h of acid gas is sent to SRU plant 9.


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CHAPTER 8

ENERGY BALANCE

8.1 Heat capacity constants for the components of sour gas

Feed gascomposition

kmol/h A B*103

C*106

D*109

H2S 27.21 34.523 -17.64 67.66 -53.24

H2 36.32 28.6105 1.0194 -0.1476 0.769

C1 59 19.249 52.113 11.973 -11.317

C2 35.7 4.1261 155.021 -81.545 16.975

C3 46.6 -4.222 306.26 -158.63 32.145

i- C4 18.9 -8.9133 419.53 233.633 51.043

n- C4 20.4 -2.451 391.82 -202.98 40.793

n-C5 1.3 -3.6266 487.4859 -258.0312 53.0488

C6 1.6 -4.4152 581.9223 -311.8584 64.9193

Total 247.03

Table 8- 1Heat capacity constants for the components of sour gas


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8.2 Energy balance across the absorber

Figure 8- 1 Energy balance across the absorber

Enthalpy of sour gas entering at 314k:Feed gas

composition

ni

kmol/h

ni *a*10 ni *b*10 ni *c*10

ni *d*10

9

H2S 27.21 939.38 -480.20 1841.20 -1448.81

H2 36.32 1039.13 37.02 -5.36 27.93

C1 59 1135.71 3074.7 706.41 -667.72

C2 35.7 193.24 6357.71 -2405.28 311.11

C3 46.6 -196.98 14271.90 -7392.23 2472.07

i- C4 18.9 -166.94 7929.19 4415.67 964.72

n- C4 20.4 -50 7993.28 -4140.96 832.17

n-C5 1.3 -4.71 398.14 -206.22 41.79

C6 1.6 -7.06 931.08 -498.97 103.87

Total 247.03 2881.77 40512.82 -7685.74 2637.13

Table 8- 2 Enthalpy of sour gas entering at 314k


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Q= 298314

(2881.77 + 40512.82*10-3

T - 7685.74*10-6

T2

+ 2637.13*10-9

T3

)dT

Q= 65.04 KW

Enthalpy of sweetened gas leaving absorber at 311k:

Feed gas

composition

nikmol/h

ni *a*101

ni *b*103 ni *c*10

6 ni *d*10

9

H2S Nil nil Nil nil nil

H2 36.32 1039.13 37.02 -5.36 27.93

C1 59 1135.71 3074.7 706.41 -667.72

C2 35.7 193.24 6357.71 -2405.28 311.11

C3 46.6 -196.98 14271.90 -7392.23 2472.07

i- C4 18.9 -166.94 7929.19 4415.67 964.72

n- C4 20.4 -50 7993.28 -4140.96 832.17

n-C5 1.3 -4.71 398.14 -206.22 41.79

C6 1.6 -7.06 931.08 -498.97 103.87

Total 219.82 1942.39 40993.02 -9526.94 4085.94

Table 8- 3 Enthalpy of sweetened gas leaving absorber at 311k

Q= 298 311

(1942.39 + 40993*10-3

T - 9526.94*10-6

T2

+ 4085.94*10-9

T3

) dT

Q= 49.32 KW


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Heat generated due to absorption:The exothermic heat of reaction @ 298.15 k for the absorption of H2S is 1047 kJ / kg.

Therefore, heat generated due to absorption = (27.21*34)*1047

= 269.06 KW--------- (1)

Heat given up in the absorber by the gas mixture:= Enthalpy of (sour gas-sweet gas)

= 52.65 - 49.32

= 15.72 KW---------- (2)

Enthalpy of Lean Amine = (170.97 * 53.35) * 3.77 * (316-298)=171.94 KW--------- (3)

Enthalpy of Rich Amine = (1) + (2) + (3)= 456.72 KW

Overall energy balance:

Enthalpy of sour gas in +

Heat generated due to absorption + = Enthalpy of sweetened gas+

Heat given off by the sour gas mixture + Enthalpy of Rich Amine

Enthalpy of Lean Amine

65.04 + 269.06 + 15.72 + 171.94 = 456.72 + 49.32 KW


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8.3 Energy balance across the solution interchanger

Figure 8- 2 Energy balance across the solution interchanger

In this exchanger, Lean amine cools down from 389 k to 353 k.

Enthalpy of Rich Amine- shell side in = 456.72 KWHeat given upas the solution cools down = 9121.25 * 3.77 * (389 - 353)

= 1299.07 KW

This much heat would be picked up by the Rich Amine entering at 340.51 k with 2% radiation

loss.

Thus heat absorbed by Rich Amine = 1299.07 * 0.98= 1273.09 KW

Enthalpy of Rich Amine- shell side out = 1273.09 + 456.72= 1729.08 KW

Enthalpy of Lean Amine- Tube side in = 869.23 KW Enthalpy of Lean Amine- Tube side out = 525.36 KW

Lean Amine

389 k

9121.25 Kg/h

Tube side in

Rich Amine

340.51 k

9121.25 Kg/h

Shell side in

353 k

459.01 k


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Overall energy balance across the Heat Exchanger:

Q in = Q out

Enthalpy of ( Lean Amine-in + Rich Amine-in)+ Heat released + Heat lost due to radiation

= Enthalpy of ( Lean Amine-out + rich Amine-out)

456.72 + 869.23 + 1299.07 + 54 = 525.36 + 1729.08 KW

Finding the temperature of Rich Amine leaving the solution interchanger.

Q = m Cp dt

1273.09 KW*3600 = 10046 * 3.85 *(T-340.51)

Thus temperature of Rich Amine is 340.51 k


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8.5 Energy balance across the stripper

Figure 8- 4 Energy balance across the stripper

Finding the temperature of the stripped out gas from stripper:

Applying Raoults law for this purpose :

ya=Pa/P and also xa=Pa/Pa

Combining the above two expressions,

xa Pa = Pa =ya P where

Pa = Partial pressure of the component A in gas

Pa = Vapour pressure of component A in liquid

xa = Mole fraction of water vapour in the gas leaving stripper

ya = Mole fraction of water in the total liquid in the stripper


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Finding ya =

= (27.125+1.70) / (27.125+1.70+27.125)

= 0.52

Therefore, Pa = 0.52*150 kPa

= 77.28 kPa

Now finding Xa =

= (1.70+ 170.97*0.65)/(170.97+27.125)

= 0.57

And now Pa = Pa/ xa

= 77.28/0.52

= 148.62 kPa

From the vapour pressures table,

Temperature of gas stream leaving the stripper=107.13C = 380k

Enthalpy of Heat in streams:

1) Enthalpy of rich amine stream =10046*3.85*(459.01-298)

= 1729.83 KW-------------- (1)

2) Enthalpy of Reflux + Make up H20 @40c

= (27.125+1.70)* 18 * (h@40C-h@25C)

= 518.85 * (168.771 - 104.77)

= 9.22 KW-------------- (2)

3) Heat supplied by the reboiler @ 2% radiation loss

= (x*0.98) KW-------------- (3)


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Enthalpy of heat out streams:

4) Heat generated due to desorption = 269.06 KW-------------- (4)

5) Enthalpy of Lean Amine stream = 9121.25 * 3.77 * (389-298)

= 869.23 KW-------------- (5)

6) Enthalpy of strip out gases at 380 k

Feed gascomposition

nikmol/h

ni *a*10 ni *b*10 ni *c*10 ni *d*10

H2S 27.21 939.38 -480.20 1841.20 -1448.81

H2O 28.91 1469.93 6160.14 -16412.52 18347.20

Total 56.12 2409.31 56810 -16412.52 18347.20

Table 8- 4 Enthalpy of strip out gases at 380 k

Q = 298380

(2409.31 + 56810*10-3

T - 16412.52*10-6

T2

+ 18347.20*10-9

T3

) dT

Q = 72.09 KW-------------- (6)

Overall energy balance across the stripper:

(1)+ (2) + (3) = (4) + (5) + (6)

1729.83 + 9.22 + (x*0.98) = 269.06 + 869.23 + 72.09

x = -539.46 KW

i.e. 539.46 KW has to be supplied by Reboiler.

For serving this purpose, a saturated steam @ 10.5 bar supplied of latent heat

= 2006 kJ/kg

Therefore , steam requirement in reboiler = 539.46 / 2006

= 0.27 kg/s

= 968.12 kg/h


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8.6 Condenser duty

Figure 8- 5 Condenser duty

Condenser duty at 2%radiation loss = Enthalpy of (Stripped gas- acid gas to SRUReflux)

Enthalpy of Acid gas to SRU at 313K:

Feed gas

composition

nikmol/h

ni *a*101

ni *b*103 ni *c*10

6 ni *d*10

9

H2S 27.21 939.38 -480.20 1841.20 -1448.81

H2O 1.70 86.44 362.24 -1073.38 1164.09

Total 28.90 1025.82 -117.96 767.82 -284.72

Table 8- 5 Enthalpy of Acid gas to SRU at 313K:

Q= 298313

(1025.82 - 117.96*10-3T + 767.82*10-6 T 2 - 284.72*10-9 T 3 )dT

Q= 4.40 KW

Enthalpy of Reflux:

= (27.21*18) ( h at40C- h at25C)

=27.21*18 (168.771-104.771)

= 8.71 KW


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Therefore condenser duty at 2%radiation loss

= (72.09 - 4.40 - 8.71) / 0.98

= 60.18 KW

Cooling water requirement:

Q = m Cp dt

60.18*3600 = m*4.1886*10

m =5.172 tonnes /h


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CHAPTER 9

HEAT EXCHANGER DESIGN

9.1 Plant data:Shell side temperature (cooling water) = 32C42C

Tube side temperature (lean amine) = 80C43C

Thermal Conductivity of lean amine = 0.37 W/ mC

Thermal Conductivity of cooling water = 0.519 W/ mC

Lean amine flow rate from stripper to absorber (tube side) = 9121.25 kg/h

Cooling water flow rate (shell side) = 3038.89 kg/h

Shell side fluid: Cooling water

Tube side fluid: Hot lean Amine

Heat load (Q) = 9121.25 * 3.77 * (80-43) / 3600

= 353.42 KW

TLMTD = ( 38-11) / ln (38/11)

= 21.780

C


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Picturistic description:

Figure 9- 1 Hot amine cooler

9.2 Properties of the fluidProperties Tube Side Shell Side

Density 1003 kg/m 995 kg/m

Specific Heat Capacity 4.18 kJ/kg K 3.77 kJ/kg K

Viscosity 0.00033 kg/ms 0.00095 kg/ms

Thermal Conductivity 0.37 W/mk 0.519 W/mk

Heat profile:

hot amine

80C 43C

42C 32C

cold water


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9.3 Calculation of P & RP = dTh / (Th1- Tc1) =10 / 48 P= R*P = 0.77

= 0.2083

R = dTh / dTc = 37/10 R= 1/R = 0.27

= 3.7

Graph for finding the correction factor :

Figure 9- 2 Graph for finding the correction factor


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To find Correction Factor, consider 1 Shell and even even Tube pass

Ft = 0.87

Tm = Ft* TLMTD

Tm = 0.87 * 21.78

Tm = 18.95C

Assuming an Overall heat transfer coefficient of

U = 500 W/m2 0

C

Provisional Area:

A = Q/U Tm

A = 353.42*103

/ 500*18.95

A = 37.3 m2

Tube description:

Outer Diameter = 20 mm

Bwg = 14

Internal Diameter = 15.8 mm

Length = 4880 mm

Allowing for tube sheet thickness take Le = 4830 mm

Pitch = 25 mm


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Area of one tube = * Do * Le

= * 0.02*4.83

=0.3033 m2

= 0.3630 m2

Number of tubes (Nt) = provisional area / area of one tube

Nt = 37.3 / 0.3033

Nt = 123 tubes

Therefore, number of tube passes is 4.

Shell description:

Inner Diameter = 387.35 mm

Number of shell side passes = 1 as theres a temperature approach only.

9.4 Film coefficient - tube side-lean aminehi Di/ k = 0.023 Re

0.8Pr

0.33(/w )

0.14

Re = Gt Di / where Gt = Mass velocity of the fluid in tube side.

Gt = mh /At where At = area of tube side per pass.

At = /4 * Di2 * Nt/ Np

= 0.0060 m2

Gt = 2.534 / 0.0060

= 422.33 kg/m2s


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Re = 422.33 * 0.0158/ 0.33 * 10-3

=20220.7921

Pr = Cp / k

= 3.77*103*0.33*10-3

/ 0.37

= 3.362

hi = 0.023 * 20220.79210.8

* 3.3620.33

(1)0.14

* 0.37 / 0.0158

= 2237.12 W/m2

C

hio = hi *Di/Do

hio = 2237.12 * 0.0158 / 0.02

hio = 1767.325 W/ m2

C

9.5 Film coefficient - Shell side-cooling waterho De/ kf = 0.36 Re

0.55pr

0.33(/w )

0.14

for triangular pitch De = 4/Do * ( 0.86 Pt2- /4 Do2 )

= 4/*0.02 * ( 0.86 0.0252- /4 0.022 )

= 0.0127 m

Gs = mc/ As

As = B c D / Pt The baffle cut is kept to be 40 % of the shell ID.

= (0.4 * 0.387) *0.005 * 0.387 / 0.025

= 0.0120 m2


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Gs = 0.8441 / 0.0120

= 70.3417 kg/ m2s

Re = Gs De/

= 70.3417 * 0.0127 / 0.95 * 10-3

= 940.3575

Pr = Cp / k

= 0.95*10-3

* 4.18*103

/ 0.519

= 6.7305

ho = 0.36 *940.35750.55

* 6.73050.33

(1)0.14

* 0.519 / 0.0127

ho =1354.93 W/m2

C

9.6 Tube wall resistanceRw = Do ln (Do/Di ) / 2 Kw

Rw = 1*10-6

W/m2

C

Considering a combined dirt factor of 0.0006 m2

C / W

9.7 Overall heat transfer coefficientUo = )-1

Uo = 524.97 W/m2

C


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9.8 Heat transfer area requiredAr = Q / Ua Tm

Ar = 353.42*103/ 524.97*18.95

Ar = 35.53 m2

% excess = Aa-Ar/Ar * 100

% excess = 37.3-35.53/ 35.53 *100

% excess = 5% so the design is safe.

9.9 Pressure drop tube side

Pt =

+ 2.5 Np v

2s

f = 0.72 Re-0.33

= 0.0273

v = Gt/

= 422.33/ 1030

= 0.4100 m/s

L = 4.88 m

Np = 4

Di = 0.0158 m

= 1


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s = 1.03

Pt = 1.460+0.8657

Pt = 2.3258 kPa

9.10 Pressure drop shell side

Ps =

f = 1.87 Re-0.2

= 1.87 * 940.3572-0.2

= 0.4755

Gs = 70.3417 kg/m2s

De = 0.0127 m

Nb =

= 4.8 / 0.01548

= 31

s = 0.995

=1

Ps = 1.147 kPa.


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9.11 Design summary The cooler is designed to be consisting of tubes of 20mm OD * 15.8 mm ID * 4880 mm

long and a shell ID of 387.35 mm of passes 4 and 1 respectively.

The overall heat transfer coefficient for the respective duty is found to be 524.97 W/m2C.

A pressure drop of 2.3258 kPa is achieved in tube side and a drop of 1.147 kPa in shellside.


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The Antoinnes equation for H2S:

log10 P =

Where pressure is in bar and temperature is in k. As calculated, the Vapour pressure of H2S is

107.720 bar.

So K =

=

= 53.860

The average K values of the components at the stripper operating pressure of 2 bar are given in

the following table. An average vapour to liquid ratio of about 0.05 is taken.

Components Avg.K S.F= VK/L

H2S 53.860 2.69

H2O 0.743 0.037

10.3 Calculating the ideal number of trays:The ideal number of trays can be calculated from the following Kremser equation

correlation graph for a stripping factor of 2.69 and 100% stripping.

So the ideal number of trays = 4

10.4

Calculating the actual number of traysAs stripping has a very low efficiency due to desorption of gases by steam

stripping from a liquid, that too in larger sized column. Due to this, the tray efficiency varies

from 12 % to 18%. Assuming an average efficiency of 13%,

The actual number of trays = 4 /0.13 = 30 trays


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10.5 Kremsers correlation graph

Figure 10- 1Kremsers correlation graph


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10.6 The Internal Diameter of the stripping columnFor finding the Internal Diameter of the column, the cross sectional area has to be calculated.

The cross sectional area of the tower will be based on the vapour and liquid flow rates at the

bottom tray.

Finding the total amount of vapour to the bottom tray:

The total vapour to the tray =27.125 + 27.125 + 1.70

= 55.95 kmol/hr

Finding the volume occupiedby the vapour thats stripped out of the column at the prevailing

conditions:

At standard Temperature and pressure

P1 =101.325 k Pa

T1 =273 k

As we know that at NTP ( thats P=101.325 kPa and T=273 k) 1 kmol of gas occupies 22.414 m3.

V1 = 55.95 * 22.414

=1254.06 m3/h

And now the prevailing conditions of pressure and temperature in the stripper are as follows :

P2 = 148.62 kPa

T2 = 380 k


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By volumetric law thats by the combination of Boyles law and Charles law,

=

V2 =

V2 =

V2 = 1190.08 m3/h

Finding the total vapour flow rate in terms of mass flow rate:

Mass flow rate = (27.125+1.70) * 18 + 27.125 *34

= 518.85 + 922.25

= 1441.1 kg/h

Finding the total amount of liquid from the bottom tray:

Total liquid from the bottom tray i.e. the lean amine flow rate

= 198.18 kmol/h

= 10046 kg/h

= 10.046 m3/h


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The following tabular column shows the total liquid and vapour flow rates to and from the

bottom tray respectively:

kmol/h kg/h m3/h

liquid flow rate 198.18 10046 10.046

vapour flow rate 55.95 1441.1 1190.08

Table 7- 2 The total liquid and vapour flow rates

Finding the vapour density and liquid density:

vapour density =

= 1.21 kg/m3

liquid density =

= 1000 kg/m3

Finding the Brown and Souder flood constant :

For the assumed tray spacing of 15 inches or 38.1 cm, of sieve trays the K factor is found from

the below given graph as K= 111 m/s ( 600 ft./s from graph)

Loading at flood = K

= 111= 3858 kg/h m

2

The design loading is to be 80 % of the loading at flooding.

Therefore, design loading = 0.8 * 3858 kg/h m2

= 3086 kg/h m2


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The cross sectional area of the column =

=

= 0.1085 m2


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Graph for finding the Brown and Souders flood constant:

Figure 10- 2 Graph for finding the Brown and Souders flood constant


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Finding the Internal Diameter of the column:

Therefore Internal Diameter of the column ID =

= 0.38 m

= 1.2 ft.

2 ft.

10.7 Calculation of column height

Height between bottom tray and top tray:

Height between bottom tray and top tray = Tray spacing (No.of trays - 1)

= 38.1 * ( 30-1)

= 11.05 m

Height between top tray and top tan:

The top tray to top tan distance is generally fixed at around 1m.

Height between bottom tray and bottom tan:

The resident time of the liquid hold up is based on the pumping capacity.

Assuming a hold up time of 2 minutes.


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Therefore, Volume resident = * 2

=

* 2

= 0.32 m3.

Height of the liquid level =

= 2.95 m say 3 m.

As an allowance of 1 m has to be given so the bottom tray to bottom tan is 4m.

So, the tan-to-tan distance = 16 m.

And now, a skirting of 5 m is also provided which makes the entire structure to rise to a

height of 21 m or 69 ft.

10.8 Design summaryThus a stripper of ID 2 ft., height 69 ft., and with 30 trays each of 15 inch spacing

is designed for the purpose.


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CHAPTER 11

COST ESTIMATION

11.1 IntroductionCost estimation is the approximation of the cost of a program, project, or operation. The

cost estimate is the product of the cost estimating process. The cost estimate has a single total

value and may have identifiable component values. A problem with a cost overrun can be

avoided with a credible, reliable, and accurate cost estimate. The cost estimation of the sour gas

treating unit , plant 12 of CPCL are as follows.

11.2 Total Capital InvestmentTotal capital investment represents the total cash investment that shareholders and debt

holders have made in a company.

Total capital investment is used in several important measurements of financial

performance, including return on invested capital, economic value added and free cash flow.

Estimation of Total Capital Investment

The total capital investment I involves the following:

The fixed capital investment in the process area, IF. The capital investment in the auxiliary

services, IA. The capital investment as working capital, IW.

i.e., I = IF + IA + IW

Estimation of Fixed Capital Investment in the process area, IF:

This is the investment in all processing equipment within the processing area. Fixed

capital investment in the process area, IF = Direct plant cost + Indirect plant cost

The approximate delivered cost of major equipments used in the sour gas treating plant-

12 of CPCL are furnished below:


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S.No.Equipment Units Cost in lakhs/unit Cost in lakhs

1 Absorption column 1 150 150

2 Stripping column 1 150 150

3 Exchanger and

cooler

2 50 100

4 Filter 1 15 15

5Pumps, accumulator etc.,

tanks,etc

2+1 25+50 75

Total 490 lakhs

Table 11- 1 Estimation of fixed capital investment

11.3 Calculation of direct cost factorA price that can be completely attributed to the production of specific goods or services.

Direct costs refer to materials, labour and expenses related to the production of a product.

S.No ItemsDirect Cost Factor

(%)

Cost in Lakhs

1 Delivered cost of major equipments 100 490.00

2 Equipment installation 25 122.5

3 Insulation 15 73.5

4 Instrumentation 20 98


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5 Piping 75 367.5

6 Land and building 30 147

7 Foundation 10 49

8 Electrical 20 98

9 Clean up 5 24.5

Total 300 1470

Table 11- 2 Calculation of direct cost factor

Direct plant cost = [(Delivered cost of major equipments) x (Total direct cost factor) / 100]

Direct plant cost = [(490 x 300) / 100]

= 1470 lakhs

11.4 Calculation of Indirect Cost factorCosts, such as depreciation or administrative expenses, are more difficult to assign to a

specific product, and therefore are considered indirect costs.

S.No Item Indirect Cost Factor Cost in Lakhs

1 Overhead Contactor 30 441

2 Engineering fee 13 191.13 Contingency 13 191.1

Total 56 823.2

Table 11- 3 Calculation of indirect cost factor

Indirect plant cost = [(Direct plant cost) x (Total indirect cost factor))/100]

= [(1470 x 56)/100)]

= 823.2 lakhs

Fixed capital investment in the process area,

IF = Direct plant cost + Indirect plant cost

= 1470+823.2

= 2293.2 lakhs


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11.5 Estimation of the capital investment in the auxiliary service, IASuch items as steam generators, fuel stations and fire protection facilities are commonly

stationed outside the process area and serve the system under consideration.

S.No Items Auxiliary services cost Cost in Lakhs

1 Auxiliary buildings 5 114.66

2 Water supply 2 45.864

3 Electric Main Sub station 2 45.864

4 Process waste system 1 22.932

5 Raw material storage 2 45.864

6 Fire protection system 1.5 34.398

7 Roads 1.0 22.932

8 Sanitary and waste disposal 0.5 11.466

9 Communication 0.5 11.466

10 Yard and fence lighting 0.5 11.466

Total 16 366.912

Table 11- 4 Estimation of the capital investment in the auxiliary service

Capital investment in the auxillary services = [(Fixed capital investment in the process area) x

(Auxiliary services cost factor) / 100]

= 366.912 lakhs

Installed cost = Fixed capital investment in the process area + Capital investment in the auxiliary

service

= 2293.2 + 366.912

= 2660.112 lakhs


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11.6 Estimation of the capital investment as working capital, IWThis is the capital invested in the form of cash to meet day-to-day operational expenses,

inventories of raw materials and products. The working capital may be assumed as 15% of the

total capital investment made in plant.

Capital investment as working capital, IW = [(2660.112) x 15) / 85]

= 39901.68/ 85

= 469.43 lakhs

11.7 Estimation of total capital investmentTotal capital investment, I = IF+ IA+ IW

= 2293.2+366.912+469.43

= 3129.54 lakhs


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CHAPTER 12

PLANT LOCATION AND LAY OUT

12.1 Plant layoutThe arrangements of equipment and facilities specified from process flow sheet

considerations is a necessary requirement for accurate preconstruction cost estimation or for

future detailed design involving piping, structural and electrical facilities. Careful attention to the

development of plot and elevation plans will point out unusual plant requirements and therefore

give reliable information on building and site costs required for precise preconstruction cost

accounting.

12.2

Factors in planning lay-outRational design must include arrangement of processing areas and handing areas

in efficient coordination and with regard to such factors as,

1. New site development or addition to a previously developed site.

2. Future expansion.

3. Economic distribution of serviceswater, process steam, power and gas.

4. Weather conditionsare they amenable to outdoor construction.

5. Safety considerationpossible hazards of fire explosion and fumes.

6. Building code requirements.

7. Waste disposal problems.

8. Sensible use of floor and elevation space.

12.3 Methods of lay-out planningTo start a detailed planning study, space requirements must be known for various

products, by-products and raw-materials as well as for process equipment. A starting or reference

point, together with a directional schematic slow pattern will enable the design engineers to

make a trial plot as shown in the figure. A number of such studies will be required before a

suitable plot and elevation plan is chosen. The various methods of Lay-out are as follows:


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Unit area concept

The basic blocks with which to build an arrangement for plot plans are often used

in the unit area concept. This method is particularly well adopted to large plan Lay-outs. Unit

areas are often delineated by means of distinct process phases and operational procedures, by

reason of contamination and by safety requirements.

Two-dimensional Lay-out

To visualize the layout problem, two dimensional scaled templates or small

cutouts of unit areas and equipment within each areas are shifted about on crosshatched scale

paper.

Figure 12- 1 Plant layout


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12.4 Scale ModelsRecent developments in the use of scale models have shown the advantage of this

method over the detailed two-dimensional method. The method is used to develop plot and

elevation plans and can also be used for piping and utilities layout with a good dimensional

accuracy. The advantage of three-dimensional models can be summarized as,

1) Optimum design selection.

2) Effective construction planning.

3) Savings in engineering design, construction, operating and maintenance cost

4) More rapid and safer training of personnel.

12.5 Equipment LayoutIn making a layout, ample space is assigned to each piece of equipment.

Accessibility is an important factor of maintenance. It is extremely poor economy to fit the

equipment layout too closely into a building. A slightly large building than appears necessary

will cost little more than one this is crowded. The extra cost will indeed be small in comparison

with the penalties that will be extracted if, in order to iron out the kinds, the building must be

expanded.

The operations that constitute a process are essentially a series of unit operationsthat may be carried on simultaneously. These include filtration, crystallization, separation, and

drying. Since these operations are repeated several times in the flow of materials it should be

possible to arrange the necessary equipments into groups of the same kinds. This sort of layout

will make possible a division of operating labour so that the one or two operators can be detailed

to tend all equipment of a like nature.

The relative levels of the several pieces of equipment and their accessories

determine their placement. Although gravity flow is usually preferable it is not altogether

necessary because liquids can be transported by blowing or by pumping and solids can be moved

by mechanical means. Access for initial construction and maintenance is a necessary part to

planning. For example, overhead equipment must have space for lowering into place, and heat

exchange equipment should be located near access areas where trucks and hoists can be placed

for pulling and replacing tube bundles. Thus space should be provided for repair and replacement


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of equipment and also for access way around doors and underground hatches. But a little regard

should be given for proper placement of each of these services, practicing good design, aids in

ease of operation, orderliness, and reduction in costs of maintenance.

Building

After a complete study of qualitative factors, the selection of the building or

buildings must be considered. A complete architect should be consulted to design a building

around the process. It is fundamental chemical engineering industries that the buildings should

be built around the process, instead of the process being made to fit buildings of conventional

design.

Waste-DisposalThroughout the chemical industry much thought should be given to the disposal

of waste liquor, fumes, dusts and gases. Ventilation, fume elimination and drainage may require

the installation of extra equipment. In the selection and the placement in the process area, such

pieces of equipment for doing the above services is also an important point. Sometimes air

conditioning of the plant is called for that may require an elaborate set up.

MaterialsHandling Equipment

Consideration of equipment for material handling is only a minor factor in most

cases of arrangement owing to the multiplicity of available material handling devices. But

where this operation is paramount in a process serious thought must be given to it.

12.6 Plant LocationThe various needs and conditions must be defined clearly before considering a

site to locate a factory. The main factor that is to be considered for deciding the location of a site

for a factory is that it should yield maximum possible benefit with minimum possible efforts.

The many number of factors that are to be taken into account to satisfy the above

statement are broadly divided into two groups.

1) Primary factors

2) Secondary factors


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Primary factors

Availability of raw materials

The location of the factory is influenced by the sources of raw materials. It is definitely

advantageous things of the factory is located near the source of raw materials.

Availability of power, Fuel

Power is an essential requirement for a factory. Therefore sufficient quantity of power should be

available for the needs of the factory without failure.

Availability of water

Water is important requirement for a factory. It is used for cooling the reactors. So, plant islocated near the rivers.

Availability of labour

Availability of required number of labour in the requisite type is also an essential consideration.

Density of population is not only the consideration. The people in surrounding areas are willing

to work the factory. This depends on the maturity of citizens, influence of local industries lively

hood, etc.

Climatic and atmosphere conditions

If the climatic and atmosphere conditions are not favourable, employees may not be willing to

work. The second consideration is the precision machines get spoiled.

a) If there is much fluctuation of temperature.

b) If the factory is within the vicinity of the sea breeze it requires a cool atmosphere. In modern

days the temperature inside the factory is regulated by scientific method.

Market facilities

Facilities for marketing the finished goods should be available. If the markets are near to factory,

a) The cost of transportation will be less.

b) The finished goods will be known in the market.


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Transport facilities

Economical conveyance a quick transport facilities should always be available whenever the raw

materials, etc. are to be transported to the factory or finished products to the markets.

Road, rail, water or air transportation facilities are required for the plant under consideration.

Secondary Factors

The following factors are also be taken into account before selecting a site for a factory.

1. Cost of land

2. Laws governing the area

3. Fire protection

4. Availability of water for the employees and factory purposes

5. Progressive attitude of the citizens of the area towards the industry


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safe trips for critical equipment. Static electricity has been cause of many fires in the oil industry.

Refineries also deal with toxic chemicals. Health monitoring is necessary for the operating

personnel who work in the areas where these toxic chemicals are handled.

13.4 LayoutLayout consideration has to be given for the inter spacing between

1) Equipment within the process block

2) Utilities

3) Tankage & Loading gantries to minimize the involvement of adjacent equipment in the event

of fire.

4) The effects of radiant heat from storage tanks & the BLEVE effects from LPG storage vessels

are to be given serious consideration.

13.5 Operating ProceduresIn many of the worst disasters that took place all over the world, it was not the

technology that was lacking rather the discipline required to follow procedures & good operating

practices. Adhering to written procedures, keeping the instruments healthy & online are some of

the essential items of a good operating discipline.

13.6 Static EquipmentCrude containing naphthenic acid, which affects transfer lines. Vessels & piping in MEA

services are susceptible to cracking. All equipment, pressure vessels, Pipe lines, tankage are to be

periodically checked for corrosion aspects for different operating conditions.

13.7 FireFire in petroleum industry can breakout due to any of the following reasons, Spark by

Electrical circuit, Hot work, Storage vessel not earthed, etc,. Hence a good fire protection

systems like Detectors, Monitors, Sprinklers, etc, should be considered inside the plant to kill the

fire if triggered.


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13.8 Health HazardHealth hazard for men can be caused due to air/water/thermal/noise pollution due to H 2S,

SO2, etc., higher than acceptable limits flying of solid catalyst dust, solid sulphur particles Acid

& Alkalis burns due to handling Leakage/Spillage of toxic chemical, such as phenol,

formaldehyde, chlorine, etc.

13.9 Important Safety Rules to Prevent High Pressure Hazards1. Pressure gauges like Bourdon type should be free of air inside the tube. The gauges

should be installed above eye level. Tube should never be allowed to corrode. Too rapid rise or

fall of pressure must be avoided. Root valve must be opened only to the required level & not

wide open.

2. Flammable gases at HP on rubbing against wall of metallic pipes causes static

electricity discharge. Good earthing is essential.

3. Bursting of rupture disc in vessels cause sudden high temperature which may ignite

flammable gases. Sufficient venting height is necessary, where it can cause no harm. Discharge

line must have sufficient diameter to carry the excess load release.

4. Direct fires vessels should not be emptied suddenly. The fluid running in slow

velocities slowly cools them.

5. Two Safety valves, one operating at slightly lower pressure is advantageous in very

high pressure equipment's.

6. HP tubing's should be anchored securely at frequent intervals to prevent whipping in

case of break.

7. Welded equipment should be heat treated to relieve locked up stresses. Expert welders

should be engaged. Too frequent radiography test of these equipment must be carried out for any

possible cracks developed during operation.

8. Oils should never be used as lubricant when oxygen is compressed.

9. Regular inspection should be carried out with experienced staff.


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13.10Environmental ControlEmission norms as stipulated by Central Pollution Control Board and other mandatory

regulations shall be satisfied for the project. Liquid effluent shall conform to MINAS standards.

All unit designers have to necessarily furnish complete effluent characteristics in theformat of Exhibit-I.


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CHAPTER 14

CONCLUSION

Firstly, the organization profile has been given and followed by that the characteristics of

crude oil is explained. The various refinery operations are also mentioned.

The Material and Energy balances have been developed to the whole of the process flow

diagram.

The cooler is designed to be consisting of tubes of 20mm OD * 15.8 mm ID * 4880 mm

long and a shell ID of 387.35 mm of passes 4 and 1 respectively. The overall heat transfer

coefficient for the respective duty is found to be 524.97 W/m2 C. A pressure drop of 2.3258

k Pa is achieved in tube side and a drop of 1.147 k Pa in shell side.

A stripper of ID 2 ft., height 69 ft., and with 30 trays each of 15 inch spacing is designed

for the purpose of regeneration of rich amine.

The safety and handling procedures for the refinery are elaborately studied and given and

the plant location and layout are clearly depicted.


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REFERENCES

1. Bhatt .B.I., Vora .S.M., Stoichiometry, 4thEdition, McGrawHill.2. Coulson & Richardsons, (1999), Chemical Engineering, Vol.1, Sixth Edition.3. Coulson & Richardsons, (2002), Chemical Engineering, Vol.2, Fifth Edition.4. Coulson & Richardsons, (2005), Chemical Engineering Design, Vol.6, Fourth Edition.5. David M.Himmelblau, Basic Principles and Calculation in Chemical Engineering

Prentic Hall of India.

6.

G.K.Roy, Solved Examples in Chemical Engineering, Khanna Publishers Delhi, 6

th

Edition.

7. Harry Silla, Chemical Process Engineering Design and Economics8. James G.Speight (2006), The Chemistry and Technology of petroleum(4thEdition).9. Jullius Scherzer and A.J.Gruia Hydrocracking Science And Technology,Marcel dekker

publications.

10.McCabe L. Warren, Smith C. Smith, Harriot Peter, Unit Operations in ChemicalEngineering, Sixth edition, 2007, McGrawHill.

11.Peters and Klaus D. Timmerhaus, Process Economics and Plant Design, McGraw HillInternational Edition.

12.Process design principles: Synthesis, analysis and evaluation by13.Robert E Meyers ,Handbook of Petroleum Refining Processes14.Robert E.Treyball, Mass Transfer Operations, Tata McGraw Hill Publishing Co.3rd

Editon.

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