design

DESIGN OF A GAS PROCESSING PLANT

Submitted By

Suman Sarkar (0802019)

Abdullah Al Mamun (0802021)

Nilay Kumar Sarker (0802031)

Sabbir Alam (0802034)

A Design submitted in partial satisfaction of the requirements for the degree of

B.Sc. in Chemical Engineering

Under the supervision of

Dr. Dil Afroza Begum

Professor

Department of Chemical Engineering

Bangladesh University of Engineering & Technology (BUET)

Dhaka-1000, Bangladeh

June 2014

Acknowledgements

We would like to express our most sincere gratitude to our supervisor, Professor Dr. Dil

Afroza Begum for her inspiration, guidance, support and drive for progress has been a great

motivation during our work. We are deeply indebted to her for her sincere effort to teach us

the art of doing design works as a group. His patience, enthusiasm, immense knowledge and

encouragements have carried us through to the end of the design. Thank you for being the

greatest mentors. This dissertation would not have been possible without her.

We also wish to thank our parents and family members. Their concern, encouragement, and

advice will always be remembered and welcome.

Sincerely,

Suman Sarkar (0802019)

Abdullah Al Mamun (0802021)

Nilay Kumar Sarker (0802031)

Sabbir Alam (0802034)

Table of Contents Definition of the Project ....................................................................................................... v

1.1 Definition of the Project ............................................................................................................... vi

1.2 Raw Material ................................................................................................................................ vi

1.3 Product Specification ................................................................................................................... vi

1.4 Utility .......................................................................................................................................... vii

Selection of Process ............................................................................................................. viii

2.1 Selection of the Process .......................................................................................... ix

2.1.1 Different method of Gas sweetening Process ....................................................................... ix

2.1.2 Different method of Gas Dehydration Process ...................................................................... x

2.2 Selection Criteria ......................................................................................................................... xi

2.3 Conclusion ................................................................................................................................... xi

Process Description ................................................................................................................ xii

3.1 Process Description .................................................................................................................... xiii

3.2 Sweetening Plant ........................................................................................................................ xiii

3.3 Dehydration Plant .......................................................................................................................xiv

3.4 Stabilizing the Condensates ........................................................................................................xiv

Design Basis .................................................................................................................................. xv

4.1 Design Basis................................................................................................................................xvi

4.2 Product Specification ..................................................................................................................xvi

4.3 Raw Material Specification ........................................................................................................ xvii

4.4 Specification of Utilities ........................................................................................................... xviii

4.5 Meteorological Condition and Soil Properties of the Plant Site ..................................................xx

Process Block Diagram ....................................................................................................... xxiii

Process Flow Diagram .......................................................................................................... xxv

Engineering Calculation: Material and Energy Balance ........................... xxvii

7.1Equipment List ......................................................................................................................... xxviii

Major Equipment Design ..................................................................................................... xxix

8.1 Glycol Absorber design (1/4 of Total Flow) .............................................................................. xxx

8.2 Sizing of Pumps .......................................................................................................................... xli

8.3 Sizing of Gas Liquid Separator ...................................................................................................xlv

8.3 Heat Exchanger ......................................................................................................................... xlvii

8.4 DEA Contractor Design Calculation ............................................................................................ lvii

Mechanical Design .................................................................................................................. lxix

P & ID .............................................................................................................................................. lxxvii

Plot Plan ................................................................................................................................... lxxxiii

14.1 Estimation of Total Capital .................................................................................................. lxxxvi

Appendix ....................................................................................................................................... xcvii

Reference ........................................................................................................................................ c

Definition of the Project

1.1 Definition of the Project

A Natural Gas processing plant with a capacity of 550 MMSCFDPipeline quality Natural gasis to

be produced using raw Natural Gas from the newly invented gas wellis to be set up at Kailashtilla,

Sylhetin Bangladesh including all off sites, auxiliaries and utilities and supporting facilities.

1.2 Raw Material

Raw Natural Gas from 3 producing gas wells (550 MMSCFD)

DEAmine Solution (32.5 USGPM)

TEG (866 USGPM)

1.3 Product Specification

Composition:

COMPONENTS MOLE FRACTION

Methane

Ethane

Propane

i-Butane

n-Butane

n-Pentane

n-Hexane

H2O

Nitrogen

CO2

0.930973

0.040237

0.000658

0.000022

0.000014

0.000007

0.000002

0.000101

0.002029

0.012562

Properties of product:

Vapour/Phase Fraction

Molar Density (kgmole/m3)

Mass Density (kg/m3)

Molar Heat Capacity (KJ/kgmole-C)

Mass Heat Capacity (KJ/kg-C)

Thermal Conductivity (W/m-k)

Viscosity (cP)

Molecular Weight

Z Factor

1.0000

3.319

57.53

36.98

2.133

4.095e-002

1.357e-002

17.34

1.000

1.4 Utility

Power

Own power generation of total Capacity of 2 MW

Where one is gas generator (1350KW) & a diesel generator (650KW).

Natural Gas

Raw Natural Gas from 3 gas wells.

Selection of Process

2.1 Selection of the Process

2.1.1 Different method of Gas sweetening Process

a) Solid bed Sweetening Process

b) Aquasorption Process (Wash water process)

c) Selexol Process

d) Chemical Absorption process

e) The Holmes-Stretford Process

2.1.1.a Solid bed Sweetening Process

The processes are based upon physical or chemical of adsorption of acid gases on

solid. Simplicity, high selectivity (only H2S is removed) and process efficiency is

independent of pressure are the main advantages of the process. Solid bed Sweetening

Process are of two type- (i) The Iron Sponge Process, (ii) Molecular sieves Process.

2.1.1.bAquasorption Process (Wash water process)

This is an effective process for application of high pressure gas with high acid gas

content and high H2S to CO2 ratio. For high H2S to CO2 ratio gases, operating and

investment can be reduced by applying this process.

2.1.1.c Selexol Process

In selexol process Dimethyl ether of polyethylene glycol (DMPEG) is used as solvent.

Solubility of the acid gas in DMPEG is directly proportional to the partial pressure of

the acid gas component. H2S gas is more soluble than CO2.. Higher hydrocarbons

have the greater solubility.

2.1.1.d Chemical Absorption process

The alkanol-amin process-

The alkanol-amine process are the most prominent and widely used process for H2S

and CO2 removal. They give offer good relatively low cost and good flexibility in

design and operation. Monoethanolamine (MEA), Diethanolamine (DMA),

Teriethanolamine (TEA) etc are most common alkanol-amin.

The hot carbonate process-

An aqueous potassium carbonate solution is the main reagent of hot carbonate

process. High concentration & temperature of potassium carbonate and high partial

pressure of CO2 gas are required for better performance. It is not applicable for gas

stream which contains only H2S.

2.1.1.e The Holmes-Stretford Process

By this process H2S can be removed as elementary carbon but CO2 content remains

unaltered. This process can be designed for large temperature & pressure range.

Lower operating cost and flexibility is the main advantage of this process.

2.1.2 Different method of Gas Dehydration Process

a) Absorption dehydration Process

b) Adsorption Process

Among these, Glycol process for Dehydration and DEAmine Process for Sweetening are

used in this process design.

2.1.2.a Absorption Process

Absorption dehydrationis a process where a liquid dessicant is used for the removal of

water vapor from a gas. Glycol, Ethylene glycol (EG), diethylene glycol (DEG),

triethylene glycol (TEG), etc are the most common absorbent used in absorption

dehydration process

2.1.2.b Adsorption Process

Adsorption dehydrationis a process where a solid desicant is used for the removal of

water vapor from a gas stream. Alumina, silica gel, alumina silica gel, Molecular

sieves are the most common adsorbent used in adsorption dehydration process.

2.2 Selection Criteria

As there are many variable in Natural Gas treating process, it is difficult to choose an

appropriate application area of the process. However among the several factors the most

significant which are to be considered are

The types and concentrations of impurities present in the gas and degree of removal

desired.

Selectivity of acid gas removal required, if any

Temperature and pressure at which the sour gas is available and at which the sweet

gas is to be delivered.

Volume of the gas to be processed and its hydrocarbon composition.

CO2 and H2S ratio in the gas.

The desirability of sulfur recovery due to environmental problems or economics

2.3 Conclusion

Economic analysis over decades with respect to Bangladesh has proved that among all

processes Glycol Dehydration and Amine sweetening process are the most proven ones. So

nowadays this process is highly adopted throughout the world and involves so many possible

variations in approach. Most of the processing plants in Bangladesh are built on these

processes and have been performing at good condition. So, in this design project Glycol

Dehydration and Amine sweetening process was adopted.

Process Description

3.1 Process Description

The process of producing Pipeline quality Natural Gas is divided into two parts.

1) Sweetening

2) Dehydration

3) Stabilizing the Condensates.

3.2 Sweetening Plant

This process flow scheme varies little, regardless of the aqueous amine solution used as the

sweetening agent. Slight modifications can appear linked to the type of amine which is

selected and to the optimization of the scheme for specific purposes.

The general process flow for an amine sweetening plant is shown in PBD. The feed gas (sour

gas) containing H2S and/or CO2 always must enter the plant through an Inlet Separator to

remove free liquids and/or entrained solids. The gas from this separator enters the bottom of

the Absorber and flows upward trough the column in intimate counter-current contact with

the aqueous amine solution (lean solution). In the column the chemical reaction between the

amine and the feed gas acid gas occurs and the amine solution absorbs the acid gas. The

chemical reaction (due to the heat of reaction between the amine and the acid gas) is

exothermic. It will raise the temperature of the gas. Treated gas (sweet gas) leaves the top of

the column and the amine solution loaded with acid gas (rich solution) leaves the bottom of

the column.

The absorber column operates at the feed gas pressure. A minimum pressure of 4/5 b.a is

required to make the process feasible and operable. There is no limitation on high pressure as

far as the process is concerned. If the feed gas be at high temperature, a gas/gas exchanger

(using the hot treated gas as heating medium) will be provided. This equipment will be

installed upstream of the inlet separator.

The rich solution from the Amine Flash Drum then passes through an Amine/Amine Heat

Exchanger. This heat exchanger serves as a heat conservation device and lowers the total heat

requirements for the process. The rich solution is heated by the regenerated solution (lean

solution) coming from the regenerator. Then the rich solution is let down to the operating

pressure of the Regenerator (generally between 1.2 and 2 b.a) also called stripper is a

fractionation column (with trays or packing) with a condenser (using water or air as cooling

medium) and a reboiler.

3.3 Dehydration Plant

Lean, water-free glycol (purity >99%) is fed to the top of an absorber (also known as a

"glycol contactor") where it is contacted with the wet natural gas stream. The glycol removes

water from the natural gas by physical absorption and is carried out the bottom of the column.

Upon exiting the absorber the glycol stream is often referred to as "rich glycol". The dry

natural gas leaves the top of the absorption column and is fed either to a pipeline system or to

a gas plant. Glycol absorbers can be either tray columns or packed columns.

After leaving the absorber, the rich glycol is fed to a flash vessel where hydrocarbon vapors

are removed and any liquid hydrocarbons are skimmed from the glycol. This step is

necessary as the absorber is typically operated at high pressure and the pressure must be

reduced before the regeneration step. Due to the composition of the rich glycol, a vapor phase

having high hydrocarbon content will form when the pressure is lowered.

After leaving the flash vessel, the rich glycol fed to the stripper (also known as a regenerator).

The glycol is thermally regenerated to remove excess water and regain the high glycol purity.

3.4 Stabilizing the Condensates

The liquid outlets from the 2 phase separators are sent to Stabilizer. Here hydrocarbon

condensates are recovered as NGL Gases from the liquid phase. The NGL gas is the

compressed by a compressor and stored into a storage tank. This NGLs are sent to the LPG

plant for further processing.

Design Basis

4.1 Design Basis

The following is the design basis that has been assumed for the Natural Gas Processing plant

producing 550 MMSCFD pipeline quality gas by using DEAmineProcess for Sweetening and

Try Ethyle Glycol process for dehydration. The process will operate 365 days /yr.

Design basis includes site conditions, utilities and climate conditions etc. which influence the

design of individual unit, equipment or facility of the overall project. Selection of the site

conditions is extremely important in determining the ultimate success of the project.

4.2 Product Specification

Composition:


Methane

Ethane

Propane

i-Butane

n-Butane

n-Pentane

n-Hexane

H2O

Nitrogen

CO2

H2S

0.930973

0.040237

0.000658

0.000022

0.000014

0.000007

0.000002

0.000101

0.002029

0.012562

0.0012

Properties of product:

Vapour/Phase Fraction

Molar Density (kgmole/m3)

Mass Density (kg/m3)

Molar Heat Capacity (KJ/kgmole-C)

Mass Heat Capacity (KJ/kg-C)

Thermal Conductivity (W/m-k)

Viscosity (cP)

Molecular Weight

Z Factor

1.0000

3.319

57.53

36.98

2.133

4.095e-002

1.357e-002

17.34

1.000

4.3 Raw Material Specification

Raw Natural Gas

Composition:


Methane

Ethane

Propane

i-Butane

n-Butane

n-Pentane

n-Hexane

H2O

Nitrogen

CO2

0.8634

0.0392

0.0088

0.0007

0.0005

0.0005

0.0003

0.0467

0.0018

0.0204

Pressure 1400 psi

Temperature 77oF

DEAmine

Composition:

Components Mole fraction

DEAmine

H2O

CO2

0.2795

0.7187

0.0018

Try Ethylene Glycol (TEG)

Composition:

Components Mole fraction

Glycol

H2O

0.995

0.005

4.4 Specification of Utilities

Electric Power Voltage 380/220 V

Frequency 50 Hz

Phase 3 phases

Process Water

Temperature 25 oC

Pressure 0.5 MPa (G)

OH

NH

OH

OH

O

O

OH

Quality:

PH 6.5 - 8.5

Total hardness 5 mg

Total basically 5 mg

Total Fe3+ 0.1 mg/l

Turbidity 5 mg/l

Circulating Cooling Water

Inlet Outlet

Temperature < 30 oC 38 oC

Pressure 0.5 MPa 0.2 MPa

Fouling factor 6x10-4 M2.hr. oC/kcal

Turbidity 50 mg/l

Chlorine < 5 ppm

PH 6.5 8

Steam

State Saturated

Pressure 7 kg/cm2

Soft Water

Conductivity < 1x10-6 Ohm-1.cm-1

4.5 Meteorological Condition and Soil Properties of the

Plant Site

Ambient Temperature

Absolute max. Temperature: 35 oC

Absolute min. temperature: 12oC

Designed max. Temperature: 40oC

Designed min. temperature: 5oC

Atmospheric Pressure

Annual average atm. Pressure: 0.11 MPa

Max. average atm. Pressure: 0.15 MPa

Min. average atm. Pressure: 0.09 MPa

Designed average atm. Pressure: 0.2 MPa

Humidity

Annual average relative humidity: 80 %

Max. monthly average relative humidity: 87 %

Min. monthly average relative humidity: 65 %

Designed relative humidity: 90 %

Wind

Wind direction: Generally wind flows from the north to the south in the winter

season and from the south to north in the summer in our country.

Wind Velocity: 45 knot (45 nautical miles per hour or 52 miles/hr)

Rain

Annual avg. rainfall: 2850 mm

Max. Avg. rainfall: 690 mm

Max. Monthly rainfall: 235 mm

Earth bearing capacity

Load bearing capacity: The soil has a bearing capacity of 0.38 ton/ft2

required piling is used.

Corrosive tendency: Non Corrosive

Climate Condition

Summer:

Wet bulb Temperature (max) 29C

Dry bulb Temperature (max) 32C

Relative humidity 84%

Wet bulb temperature (min) 22C

Dry bulb Temperature (min) 25C

Dew point 21C


Winter:

Wet bulb Temperature (max) 22C

Dry bulb Temperature (max) 26C

Dew point 20C


Wet bulb temperature (min) 11OC

Dry bulb Temperature (min) 9C

Dew point 9C


Other Information

Natural catastrophe: A possibility of storm in the months of April-May.

Value of by-products: No economically feasible by-products are

obtained.

Process Block Diagram

Process Flow Diagram

Engineering Calculation:

Material and Energy Balance

7.1Equipment List

Table 7.1: Table of Equipment List

Equipment Name Quantity Designation in the HYSYS

silulation

Compressor 1 K-100

Heat Exchanger 2 E-100, E-102

Adsorber 2 T-100, T-102

Separator 3 V-100, V-101, V-104

Pump 2 P-100, P-101

Storage Tank 1 V-106

Stripper Column 2 V-102, V-103

Stabilizer Column 1 V-105

Major Equipment Design

8.1 Glycol Absorber design (1/4 of Total Flow) TOP Bottom

Vapor rate, ft3/sec

Liquid rate, ft3/sec

Vapor Density, lb/ft3

Liquid Density, lb/ft3

Surface tension, dynes/cm

4.1

0.49

3.6

70.31

45.87

3.7

0.84

3.92

52.12

22.1

Assumptions:

Clean service, no fouling or suspended material

Tray spacing is to be close as possible, because vertical installation space is a

premium.

All equations and data used here refer to the book A P P L I E D PROCESS D E S I G N

FOR CHEMlCAlAND PETROCHEMICA1 PlANTS, Volume 2, Third Edition by Ernest E.

Ludwig.

Calculation of Tower Diameter Souders-Brown method

W = C(v)[(l) (v)]

At the Surface Tension 22.1,

From Figure 8-82, C = 280 for 12-inch tray spacing.

In this case Vapor rates are close and vdoes not change much from bottom to top of tower.

W = 280 [3.92(52.12 3.92)]= 3848.796 lbs/hr (ft2)

= 3848.796

3600 2.872= 0.3722 ft3/sec (ft2)

Tower cross-section area, A =

=

4.1

0.3722 ft2 = 11.02 ft2

Diameter, D = 4 3.1416 = 3.75 ft

Using Hunt equation:

Assuming Height clear liquid in bubbling zone on tray, hc = hw + how = 3.5 in.

S' = St 2.5 hc = 12 2.5 3.5 = 3.25 in.

At surface tension = 22.1 dynes/cm,

For Liquid entrained, ew = 5% = 0.05

From Figure 8-121, allowable tower velocity = 1.5 ft/sec

Required tower area = 3.7

1.5= 2.47 ft2 (bottom, largest)

Diameter = 4 1.23 3.1416 = 1.78 ft

Selecting tower diameter = 4.0 ft

Tray Layout Based on 4.0 ftDiameter Tower

Using a segmental downcomer on a cross-flow tray.

From the residence time in downcomers for bubble cap trays and at the tray spacing of 12

inches, selecting an allowable liquid velocity of .4 ft/sec.

Downcomer area = . 84 .4= 2.8 ft2

Total tower area =3.1416 4.02 4= 12.57 ft2

Percent of tower area = 2.8

12.57 100 % = 22.3%

Using Figure 8-100 for segmental downcomers, at 22.3% downcomer area, the weir length is

89% of tower diameter.

Weir length, lw= 0.89 4.0 = 3.56 ft

Since standard details for fabrication are already available for a tray with a 19.5-in. weir in a

30-in. tower (65% of dia.), try this asfirst tray examined. This is 6.8% of tower cross

sectional area.

Downcomer area = 0.068 12.57 = 0.86 ft2.

Hole Size

Assuming 3/16-in. dia. 3/5-in. pitch

This is spacing of 2.66do, and is asclose as good design would suggest. Using 1/2-in. tray

thickness.

Ratio,=do

=

3/16

3/5 = 0.3125

Percent hole area = 13% (of perforation area only) as shown in Figure 8-143.

Minimum Hole Velocity: Weeping

Assuming vo (v)1/2= 13

Assuming Submergence = 4 in. = hsl= hdl (neglecting /2)

Dry Tray Pressure drop, hdt = 0.003(vo2 v)

water

L (1 2)/Co2

The hole dia

tray thickness ratio=

3/16

1/2 = 1.5

From Figure 8-126, = 0.2

From Figure 8-128, orifice coefficient, Co= 0.78

hdt = 0.00313264.2(1.22)

52.12.782= 1 in. liquid

Effective head

For hsl= 4, Fs

For hsl= ,Fs14 and hsl= 1.25 inch

Total wet tray pressure drop, ht= 5.4 + 1.35= 6.75 inches liquid

Liquid Back-up or Height in Downcomer Hd= ht+ (hw+ how) + + hd

Hd= 6.75 + 1.25 + 0 + 0 (assuming and hd to be confirmed) Hd= 8.0 in. liquid

The limit on Hd for flooding is St/2= 12/2= 6 in.

Therefore Fs= 30 appears to be a good assumption.

Design Hole Velocity

A velocity represented by Fs factor between minimum and maximum limits is to be selected.

30 >Design >25

A median value of Fs= 27.5 is selected, because freedom to operate above and below the

design value is preferred in this case.

Design Basis

Fs= 27.5

1. Weir Height selection: hw= 3.0 inch

2. Weir Length selection:For good design, ratio of weir length to tower diameter is 0.5

to 1.0

Selected weir length lw = 4.0 ft.

Weir length to tower diameter ratio is 1.0

()

2.5 =

.84448.8

4.02.5 = 11.78

From figure-8-105, correction factor: Fw: 1.0

Francis Formula , how = .092 ()

2/3 = .092

.84 448.8

4.0

2/3

in = 1.9 in

3. Submergence: hsl= (Fw) (hw) + how= (1) (3.0) + 1.9 = 4.9 in. liquid

4. Downcomer pressure loss:

Clearance between bottom of downcomer and plate = 3-in. max. Underflow area = (9.5 in.) (3

in.)/144 = 0.2 ft2. Because this is less than the downflow area, it must be used for pressure

drop determination. No inlet weir used on this design.

hdu= 0.56

Lg

449AD

2 = 0.56[

.84449

449 0.2]

2 = 9.88 in liquid

5. Dry tray pressure drop:

hdt = 0.00327.5264.2(1.12)

52.12.782= 4.55 in. liquid

6. Effective head:hsl= 4.9 in.

he= 2.55 in. liquid for Fs>14, Figure 8-130

7. Total wet tray pressure drop:

8.

ht= 4.55 + 2.55 = 7.10 in. liquid

9. Total tower pressure drop for 19 trays:

(tower)= 7.119(70.31+52.12)

17282 = 4.78 psi

which is not greater than 5 psi. So, it is satisfactory.

10. Number of holes required:

Assumed Hole size = 3/16 in.

Hole spacing or pitch = 3/5 in.

From Figure 8-144, Holes/(in.2 plate area) = 3.25

Area of a 3/16 -in. hole = 0.0276 in.2

At Design condition: Fs = 27.5

Top Vapor Velocity, vo,Top =

() =

27.5

3.6 = 14.5 ft/sec

Bottom Vapor Velocity, vo,Bottom = =

() =

27.5

3.92= 13.9 ft/sec

Number of holes required = Vapor Flow rate 144

vo hole area

at Top = 4.1144

14.5.0276= 1476

at Bottom = 3.7144

13.9.0276= 1389

At Flooding condition: Fs = 30


() =

30

3.6 = 15.82 ft/sec


() =

30

3.92= 15.16 ft/sec


vo hole area

at Top = 4.1144

15.82.0276= 1353

at Bottom = 3.7144

15.16.0276= 1274

At Weeping condition: Fs = 25


() =

25

3.6 = 13.18 ft/sec


() =

25

3.92= 12.63 ft/sec


vo hole area

at Top = 4.1144

13.18.0276= 1624

at Bottom = 3.7144

12.63.0276= 1529

11. Tower Height:

Height of the column

H = (N-1) + Ht + Hb + Htr

Total tray thickness ,Htr=

12 ft =

190.5

12 ft = 0.8 ft

Space at the top typically an additional 5 to 10 ft is needed to allow for disengaging space.

The bottom of the tower must be tall enough to serve as a liquid reservoir. At bottom portion,

top of both inlet and outlet pipes should be at least 400 to 450 mm below the bottom tray.

Additional top height, Ht= 5ft

Additional bottom height, Hb= 2.2ft

Number of tray, N= 19

H = [{(19-1) 12}/12 + 5 + 2.2 + 0.8]ft

= 26 ft

12. Mechanical tray layout details:

A total of 4-in. on diameter for extension of tray ring-type support into the tower. This

reduces available tray area.

Assuming 6.5-in. clearance (no holes) between inlet downcomer and first row of holes. The

6.5 in. could be reduced to 4 in. minimum if an inlet weir were used.

Assuming 4-in. clearance (no holes) between outlet weir and adjacent row of holes.

Downcomer width = .12 4 = 0.56 ft = 5.76 in (From Figures 8-100 at standard 65% weir

length downcomer width is 12% of Tower diameter).

Area determinations:

Area of segment of circle (2) with chord AD:

Diameter circle (2) = 48 - 4 = 44 in.

Height of chord = 44

2 (

48

2 5.76 6.5)= 10.26 in.

Chord height/circle dia. H/D =10.26

44 = 0.23

Referring to Perrys Handbook, (pg. 32, 3rd Ed.)

Area = 0.133 442= 258 in2

Area of segment of circle (2) width chord BC:


2 (

48

2 5.76 4)== 7.76 in

H/D = 7.76

44= 0.176

Area = 0.08 442= 155in2

Area of circle (2) = 442= 6028 in.2

Area available for holes = 6028 (258 + 155) = 5615 in.2

Area required for holes = 1624

3.25holes/in2= 500 in2.

Number of holes = 1624

Area = 5615 in.2

Other Mechanical Designs:

Maximum Operating Temperature = 38oC = 311k

Maximum Operating Pressure = 1250 psi = 88 kg/cm2

a) Shell Thickness:

Material = SA-516-70 Carbon Steel

Specific Gravity, s = 7.7

Maximum allowable stress, fs = 1300 kg/cm2

Welding efficiency, j = 0.9

Design Pressure, PD = 1.5 Operating Pressure = 1.5 88 = 132 kg/cm2

Internal Shell or column Dia, D = 4 ft = 1.2192 m = 1219.2 mm

Here, C = corrosion Allowance = 3 mm

Thickness, ts= PDD

2 fs j PD + C =

1321219.2

2 1300 .9 132 + 3 = 73mm

b) Tray:

Thickness = 1/2 inch = 0.5 inch


Sieve Tray, hole size = 3/16 inch, pich = 3/5 inch, Triangular.

Tray Spacing = 12 inch

c) Support for trays:

A total of 4-in. on diameter for extension of tray ring-type support into the tower.

Material: Carbon steel.

d) Downcomer and Weir:

Material = Carbon steel.

e) Support for Column:

Skirt type, Height = 6.6 ft, Material = Carbon steel.

f) Heads:

Material = SA-516-70 Carbon steel.

Permissible tensile stress = 1300 kg/cm2

Thickness of Head,

th = PDRcW

2 fs j

=1321219.21.54

2 1300 .9

= 106 mm

= 4.042 inch

Rc = Crown Radius =1219.2

W = Stress intensification factor = 1

4

3 + RcRl

= 1.54 mm

Where, Rl = Knuckle radius = 0.1 Rc

g) Insulation and Coating:

Insulating material: Polyurethane

Insulating width: 4 inch

Coating material: Zinc Sulfate

Coating width: 0.06 inch

8.2 Sizing of Pumps

Pump (P-100) of Dehydration plant

Given Density, = 972kg/m3 = 60.68 lb/ft3

Viscosity, = 17.48 cP = 11.746 10-3lb/sec.ft

Temperature = 62.370 F

Mass flow rate =14.6487kg/hr = 7110.2 kg/hr = 1.975 kg/sec = 4.35 lb/sec

Volumetric flow rate q = 072.068.60

35.4

m ft3/sec

Di opt = 3.9 q0.45 0.13 [Peters and Timmerhaus]

= 3.9 0.0720.45 60.680.13

= 2.035 inch

Let us assume, Nominal diameter = 2 inch and Schedule number = 40

Outer Diameter, OD = 2.38 inch = 0.198 ft

InnerDia, ID = 2.067 inch = 0.172 ft[From table D-13 (Peters and Timmerhaus)]

Area, A = ID2/4 = (0.172)2/4 = 0.0233 ft2

v = q/A = 0.072

0.0233=3.092ft/sec

Re = Dv

=

0.1723.0960.68

11.746 103 = 2814.15

For commercial steel /D = 0.000885 for 2.067 inch dia [from Ernest E. Ludwig (vol-1)

page- 68]

From Moody diagram (Franzini, page -226)

Friction factor f = 0.0195

Loss calculation

1. For 90 elbow(2 pcs)

F1= 2k (v2/2g)

= 20.233.09^2

232.2

=0.0682ft

2. For Tee (1pcs)

F2=k (v2/2g)

= 1.83.09^2

232.2

=0.267ft

3. For Gate valve (2 pcs)

F3 = 2k (v2/2g)

=20.193.09^2

232.2

= 0.0563 ft

4. Entrance loss

F4 = k (v2/2g)

= 0.53.09^2

232.2

= 0.0741ft

5. Exit loss

F5 = k (v2/2g)

= 13.09^2

232.2

= 0.1482ft

6. Friction loss

F6 = f(L/D)(v2/2g)

= 0.0195 110

0.172

3.09^2

232.2

=1.83ft

Total Loss, hL= F = (0.0682+ 0.267+ 0.0563 + 0.0741+ 0.1482+ 1.83)ft

= 2.4438 ft

Energy Calculation:

Using Bernoullis equation

P1/ + Z1 +V12/2g +hA -hL=P2/ +Z2 +V22/2g

Given that

P1= 20 KPa = 2.9 psi

P2= 9020KPa = 1308.6 psi

Z2-Z1= 40ft

V1=V2=3.09ft/sec

= 60% (assume)

hA=P2 P1

+ (Z2-Z1)+ (V2

2 - V12)/2g + hL

= 1308.6 2.9

60.68 32.2 + 40 + 0 + 2.4438

=43.112ft

P = hq/(500). = (60.6832.20.07243.112)/(5500.6)

= 17.1hp

Assuming over design, shaft power = 17.11.1= 18.8hp

We can select 20 HP centrifugal pump.

So straight centrifugal pump, model DUGL12A-20 was selected.

Specification of pump:

Material: cast iron

RPM: 1750

Flow capacity: 770GPM (max)

Impeller size:

Weight: 255 lbs

Price: $ 4662.00

8.3 Sizing of Gas Liquid Separator

Given,

Vapor density, v= 4.854 lbm/ft3

Liquid density, L = 55.703 lbm/ft3

Molecular weight of vapor, MWvap = 17.766 lbm/ lbmole

Vapor flow rate, V = 18.37 lbmole/sec

k = 0.40for up-flow [Ludwig, vol. 1, Page-256]

Design velocity

VD = kvl vv

= 0.40 ( 55.703 4.854

4.854)^0.5

= 1.3 ft/sec

& Actual flow rate

ACFS = V MWvap

v

= 67.23 ft3/sec

Flow area (cross sectional),

Ac=ACFS

VD ft2

= 67.23

1.3= 51.72 ft2

Inner Diameter, ID = (4 Ac /) 0.5 = (4 51.72 /3.14)0.5= 8.12 ft

For vertical separator, (h/D) min = 5

Minimum height, h = 5 8.12 = 40.6 ft

Thickness calculation

Here, operating pressure, P = 800 psi

Design pressure, PD= 1.5 P = 1.5 800 = 1200 psi

Thickness, ts = PD D

2 j0.2PD + Cs

Where

PD = Design pressure

fs= Allowable working stress =1600 psi [ Low-alloy steel for resistance to H2 & He ]

E= Joint efficiency (dimensionless) = 80%

Cs=corrosion allowance= 0.12 inch

ts = 1200 8.12 12

2 1600 0.8 0.2 1200 + 0.12 = 39.6 inch = 3.3 ft

(Ref: Peter and Timmerhaus, Chap 14, table 3 and 4, page: 537, 538 and 721)

8.3 Heat Exchanger

Thermal Design of Heat Exchanger (E-100)

Tube side (Cold fluid)

Inlet fluid: Glycol to be preheated

Outlet fluid: Preheated glycol to absorber

Inlet temperature: 67.59o F

Outlet temperature: 770F

Tavg = 72.2950F

Density (by interpolation) =70.58 lb/ ft3

Viscosity = 22.52 cP=54.5 lb/ft-hr

Thermal conductivity K= 0.192 Btu/hrft0F

Heat capacity Cp =0.5891 Btu/lb0F

Flow rate, Wg = 4.88 x 105 lb/hr

q = WgCpt = (4.88 x 105 x 0.5891 x 9.41) Btu/hr

= 2.7x 106 Btu/hr

Shell side (Hot fluid)

Inlet fluid: Dry gas

Outlet fluid: Sales gas

Inlet temperature: 99.880 F

Outlet temperature: 94.860F

Tavg = 97.370F

Density (by interpolation) = 3.601 b/ft3

Viscosity = 0.01362 cP=0.0333 lb/ft-hr

Thermal conductivity K = 0.023875 Btu/hrft0F

Heat capacity Cp = 0.5105 Btu/lb0F

Flow rate,Wi = 1.05 x 106 lb/hr

q = WiCpt = (1.05 x 106 x 0.5105 x 5.02) Btu/hr

= 2.7x 106 Btu/hr

Tube side calculation

LMTD calculation:

(First pass counter-current flow)

Hot Cold Difference

Higher 99.88 (T2) 67.59 (t1) 32.29

Lower 94.86 (T1) 77 (t2) 17.86

Difference 5.02 9.41 14.43

LMTD =(22)(11)

ln ((22)

11)

R =21

21 S=

21

T2t1

(LMTD)' = 4.39

ln (22.88/27.27)R =

5.02

9.41 S =

9.41

99.8867.59

= 25.01oF = 0.53 = 0.2914

From Fig 18 (kern) Page-828 for 1-Shell pass & 2 or more Tube pass heat exchanger

FT = 0.95 which is satisfactory..

Corrected LMTD = FT x (LMTD)'

= 0.95 25.01

= 24.510F

Calculation of heat transfer area and tube numbers

Q = Uassume Arequired LMTD

For light - heavy organics

Overall U = 10-40 Btu/hr ft20F

Assume, Uassume = 35 Btu/hr ft20F

A = 2

6

4.314751.2435

107.2ft

LMTDU

Q

assumee

Iteration- 1

The first iteration is started assuming 1 shell pass & 2 tube pass(np).

Fixed tube plate

1in. OD(d0), 14 BWG, 11

4 in. square pitch, ID (di) 0.834 in.

Assuming tube length (Lt) is 20 ft.

So, no of tubes, nt =

d0Lt= 562

Nearest count from table 9(kern) Page-841 is 574

Now, Reynolds no, Re =4wgnp

dint= 892 < 104

Iteration- 2

2nd iteration is started assuming 1 shell pass & 4 tube pass(np).

Fixed tube plate

1in. OD(d0), 14 BWG, 11




d0Lt= 561



dint= 1100 < 104

Iteration- 3

3rd iteration is started assuming 1 shell pass & 4 tube pass(np).

Fixed tube plate

1.25 in. OD(d0), 16 BWG, 19




d0Lt= 259



dint= 5647 < 104

Iteration- 4

4th iteration is started assuming 1 shell pass & 6 tube pass(np).

Fixed tube plate

1.5 in. OD(d0), 16 BWG, 17




d0Lt= 178



dint= 16560 > 104

Therefore, our final tube no 182 & Reynolds no (Re) is 16560 selected shell ID (Ds) is 35 in.

Now,

Nu =

k

hidi= 0.027 Re0.8

3

1

K

Cp(

)0.14

So, hi = 175.74 Btu/hrft2 0F

hio = hi x

=151.92 Btu/hrft2 0F

Shell side calculation

Assumption:

25 % cut segmental baffles

Baffles spacing, B = 0.5 Ds = 17.5 in = 1.46 ft (half of the shell ID is selected)

Pitch, Pt = 17

8 in. square pitch = 0.15625 ft

Clearance, C = Pt do = 0.03125 ft

Shell ID, Ds = 35 in = 2.92 ft

So, flows area, as =T

shell

P

BCID

= 0.8526 ft2

Mass velocity, Gs = Wi / as = 1.23 x 106 lb/hrft2

Equivalent dia, De =

o

T

d

dP

4/4 20

= 0.124 ft

Reynolds no, Re = .

= 465000

So,

Nu =

k

hoDe= 0.36 Re0.55

3

1

K

Cp(

)0.14

So, ho = 184 Btu/hrft2 0F

Clean overall heat transfer coefficient (Uc) calculation

Uc =(1

+ +

2+

1

+

Rd)-1 = 42.3 Btu/hrft20F

Over design calculation:

% of over design =

= 20.86 % < 30 %

So, design is accepted.

Dirt factor calculation

Clean overall heat transfer coefficient, Uc = 42.3 Btu/hrft20F

Assumed overall Coefficient, Ud = 35 Btu/hrft20F

Dirt factor, Rd = Dc

Dc

UU

UU

= 0.00098

Because for natural gas Rd, allowable=0.001 (Ref: Mechanical Design of Process System (vol-

2))

Therefore, Rd is acceptable.

Pressure drop calculation

Tube side calculation

Flow area, at= passesofNo

tubeperareaflowtubesofNo

= 6

01021.0182

= 0.309 ft2

Mass velocity,Gt = Wg/at = 1.6 104lb/hrft2

For Ret = 16560

From figure 26, tube side friction factors correlation (Kern) Page-836

Friction factor, f = f = 0.288 ft2/ft2

Assume, t =1

So, frictional pressure drop, PsitSGdi

nLtGfP

ptf 22.2

...1022.5

..10

.2

Return loss rP = 1.334 10-13 (2np 1.5)tG2.

= 0.34 Psi

So, total pressure drop in tube side, TP = fP + rP = 2.56 Psi < 8.7 psi which is acceptable.

Shell side calculation

Pitch, Pt = 17

8 in. square pitch = 0.15625 ft

Clearance, C = Pt do = 0.03125 ft

Shell ID, Ds = 35 in = 2.92 ft

Flows area, as = 0.8526 ft2

Mass velocity, Gs= 1.23 x 106 lb/hrft2

Reynolds no, Re = 465000

No of baffles, nb=

= 30.82 ~ 31

From figure 26, tube side friction factors correlation (Kern) Page-836

Friction factor, f = 0.1296 ft2/ft2

s =1(assuming)

So, pressure drop = sSGDe

nDssGfPs

b

...1022.5

)1.(..10

2

= 9.38 Psi < 14.5 Psi, which is acceptable.

Mechanical design

Tube side properties:

Materials: Stainless Steel

No. of pass: 6

Shell side properties:

Materials: Stainless Steel

No. of shell: 1

Shell dia, Ds = 35 in

Working pressure, P = 1247 Psi

Design pressure, Dp = Wp 1.1 = 1371.7 Psi

Permissible working pressure, f = 11000 Psi


Shell thickness, ts= Pjf

DsP

..2

.

= 2.088 in.

Corrosion allowance =1/8 inch= 3.175 mm

Shell thickness including corrosion allowance = (2.088+1/8)

= 2.205 in.

Nozzles

Take inlet and outlet nozzles as 100mm diameter.

Vent nozzle = 25mm diameter

Drain nozzle = 25mm diameter

Relief Valve = 50 mm diameter.

Nozzle thickness = [ P x Di ] / { 2 f J - P }

= 0.68 mm

Minimum nozzle thickness is 6mm and 8mm is chosen which includes

thecorrosionallowance.

Transverse baffles

Number of Baffles = 31

Baffle cut = 25%

Baffle thickness = 6mm (standard)

Flange design

Flange is ring type with plain face.

Flange material: IS 2004-1962 Class 2 Carbon Steel

Bolting steel: 5% Chromium, Molybdenum Steel

Gasket Material: Asbestos

Shell OD = 37.205 in

Shell Thickness = 2.205 in

Shell ID = 35 in

Allowable stress for flange material = 100 N/mm2

Allowable stress of bolting material = 138 N/mm2

Using matche.com we have estimated cost for this heat exchanger is $158900.

8.4 DEA Contractor Design Calculation

TOP Bottom

Vapor rate, ft3/sec

Liquid rate, ft3/sec

Vapor Density, lb/ft3

Liquid Density, lb/ft3

Surface tension, dynes/cm

4.1

0.49

3.65

70.20

45.60

3.73

0.85

3.96

52.00

22.00

Assumptions:

Clean service,

No fouling or suspended material,

Tray spacing is to be close as possible, because vertical installation space is a

premium.

All equations and data used here refer to the book A P P L I E D PROCESS D E S I G N

FOR CHEMlCAlANDPETROCHEMICA1 PlANTS, Volume 2, Third Edition by Ernest E.

Ludwig.

Calculation of Tower Diameter Souders-Brown method

W = C(v)[(l) (v)]

At the Surface Tension 22,

From Figure 8-82, C = 280 for 12-inch tray spacing.

In this case Vapor rates are close and vdoes not change much from bottom to top of tower.

W = 280[3.96(52.00 3.96)]= 3862 lbs/hr (ft2)

= 3862

3600 2.872= 0.3735 ft3/sec (ft2)

Tower cross-section area, A =

=

4.1

0.3735 ft2= 10.98 ft2

Diameter, D = 4 3.1416 = 3.73 ft

Using Hunt equation:

Assuming Height clear liquid in bubbling zone on tray, hc = hw + how = 3 in.

S' = St 2.5 hc = 12 2.5 3 = 4.5 in.

At surface tension = 22.0 dynes/cm,

For Liquid entrained, ew = 5% = 0.05

From Figure 8-121, allowable tower velocity = 1.5ft/sec

Required tower area = 3.73

1.5= 2.49 ft2 (bottom, largest)

Diameter = 4 2.49 3.1416 = 1.78 ft

Selecting tower diameter = 4.0 ft to meet the requirement of both bottom and top

diameter of the absorber.

Tray Layout Based on 4.0ft Diameter Tower

Using a segmental downcomer on a cross-flow tray.

From the residence time in downcomers for bubble cap trays and at the tray spacing of 12

inches, selecting an allowable liquid velocity of 0.3ft/sec.

Downcomer area = 0.85 0.3= 2.83 ft2

Total tower area =3.1416 4.02 4= 12.57ft2

Percent of tower area = 2.83

12.57 100 % = 22.5%

Using Figure 8-100 for segmental downcomers, at 22.5% downcomer area, the weir length is

89% of tower diameter.

Weir length, lw= 0.894 = 3.56 ft

Since standard details for fabrication are already available for a tray with a 19.5-in. weir in a

30-in. tower (65% of dia.), try this asfirst tray examined. This is 6.8% of tower cross

sectional area.

Downcomer area = 0.068 12.57= 0.86 ft2.

Hole Size

Assuming 3/16-in. dia. 1/2-in. pitch

This is spacing of 2.66do, and is asclose as good design would suggest. Use 1/8 in. tray

thickness.

Ratio,=do

=

3/16

1/2 = 0.375

C= 2.66do

Percent hole area = 12.8% (of perforation area only) as shown in Figure 8-143.

Minimum Hole Velocity: Weeping

Assuming Fs= vo (v)1/2= 13

Assuming Submergence = 3 in. = hsl= hdl (neglecting /2)


water

L (1 2)/Co2

The hole dia


3/16

1/8= 1.5

From Figure 8-126, = 0.128


hdt = 0.00313264.2(1.1282)

52.782= 1.01 in. liquid

Effective head

For hsl= 3, Fs


hdt = 0.00320264.2(1.1282)

52.782= 2.4 in. liquid

New Effective head

For hsl= 2.25, Fs>14

Reading from Figure 8-130; effective head, he= 1.65 in. liquid

New Total Wet Tray Pressure Drop

ht= hdt + he = 2.4 +1.65 = 4.05 in. liquid

Using Figure 8-131 Curve A, and ht= 4.05 in. liquid

Reading Weep point velocity = 21.6= vom (v)1/2

Because vom (v)1/2= 21.6 is greater than the assumed value of 20, the 20 cannot be used.

Assuming new Fs = vo (v)1/2= 25

Assuming Submergence = 1.25-in. = hsl= hdl (neglecting /2)


water

L (1 2)/Co2

The hole dia


3/16

1/2= 1.5

From Figure 8-126, = 0.128


hdt = 0.00325264.2(1.1282)

52.782= 3.74 in. liquid

New Effective head

For hsl=1.25,Fs>14

Reading from Figure 8-130; effective head, he= 1.35 in. liquid

New Total Wet Tray Pressure Drop

ht= hdt + he = 3.74 +1.35 = 5.09 in. liquid

Using Figure 8-131 Curve A, and ht= 5.09 in. liquid

Reading Weep point velocity = 24.4= vom (v)1/2

Because vom (v)1/2= 24.4 is smaller than the assumed value of 25, the 25.0 is accepted value.

Maximum Hole Velocity at Flood Conditions

Assume F, = vo (v)1/2= 28 max.

Submergence = 1.25-in. = hsl= hdl (neglecting /2)

Dry Tray Pressure drop, hdt = 0.00328264.2(1.1282)

52.782= 4.65 in. liquid

Effective head, he= 1.35 in. liquid, for Fs>14 and hsl= 1.25 inch

Total wet tray pressure drop, ht= 4.65 + 1.35= 6 inches liquid

Liquid Back-up or Height in Downcomer Hd= ht+ (hw+ how) + + hd

Hd= 6 + 1.25+ 0 + 0 (assuming and hd to be confirmed) Hd= 7.25 in. liquid

The limit on Hd for flooding is St/2= 12/2= 6 in.

Therefore Fs= 28 appears to be a good assumption.

Design Hole Velocity

A velocity represented by Fs factor between minimum and maximum limits is to be selected.

28>Design >25

A median value of Fs= 26.5 is selected, because freedom to operate above and below the

design value is preferred in this case.

Design Basis

Fs= 26.5

13. Weir Height selection: hw= 3 inch

14. Weir Length selection:For good design, ratio of weir length to tower diameter is 0.5

to 1.0

Selected weir length lw = 4.0 ft.

Weir length to tower diameter ratio is 1.0

()

2.5 =

.85448.8

4.02.5 = 11.79

From figure-8-105, correction factor: Fw: 1.0

Francis Formula, how = .092 ()

2/3

= .092 .85 448.8

4.0

2/3

in = .92 in

15. Submergence: hsl= (Fw) (hw) + how= (1) (3) + 1.92= 4.92 in. liquid

16. Downcomer pressure loss:

Clearance between bottom of downcomer and plate = 3-in. max. Underflow area =

(43.14161.52 in2.)/144 = 0.2ft2. Because this is less than the down flow area, it must be

used for pressure drop determination. No inlet weir used on this design.

hdu= 0.56

Lg

449AD

2 = 0.56[

.85448.8

449 0.2]

2 = 10.1 in liquid

17. Dry tray pressure drop:

hdt = 0.00326.5264.2(1.1282)

52.782= 4.25in. liquid

18. Effective head:

hsl= 4.92 in.

he= 2.55 in. liquid for Fs>14, Figure 8-130

19. Total wet tray pressure drop:

ht= 4.25+ 2.55= 6.8 in. liquid

20. Total tower pressure drop for 20 trays:

(Tower)=6.820(70.2+52)

17282= 4.8 psi

4.8 psi is not greater than 5 psi. So, it is satisfactory.

21. Number of holes required:

Assumed Hole size = 3/16 in.

Hole spacing or pitch =1/2 in.

From Figure 8-144, Holes/(in.2 plate area) = 5

Area of a 3/16 -in. hole = 0.0276 in.2

At Design condition: Fs = 26.5


() =

26.5

3.65 = 13.87 ft/sec


() =

26.5

3.96= 13.32 ft/sec


vo hole area

at Top = 4.1144

13.87.0276= 1542

at Bottom = 3.73144

13.32.0276= 1461

At Flooding condition: Fs = 28


() =

28

3.65 = 14.66 ft/sec


() =

28

3.96= 14.07 ft/sec


vo hole area

at Top = 4.1144

14.66.0276= 1460

at Bottom = 3.73144

14.07.0276= 1383

At Weeping condition: Fs = 25


() =

25

3.65 = 13.09 ft/sec


() =

25

3.96= 12.56ft/sec


vo hole area

at Top = 4.1144

13.09.0276= 1634

at Bottom = 3.73144

12.56.0276= 1550

22. Mechanical tray layout details:

A total of 5-in. on diameter for extension of tray ring-type support into the tower. This

reduces available tray area.

Assuming 8-in. clearance (no holes) between inlet downcomer and first row of holes. The 8

in. could be reduced to 5 in. minimum if an inlet weir were used.

Assuming 5-in. clearance (no holes) between outlet weir and adjacent row of holes.

Downcomer width = .12 4= 0.56 ft = 5.76 in (From Figures 8-100 at standard 65% weir

length downcomer width is 12% of Tower diameter).

Area determinations:

Area of segment of circle (2) with chord AD:

Diameter circle (2) = 48 - 5 = 43 in.


2 (

48

2 5.76 8)= 11.26 in.

Chord height/circle dia. H/D =11.26

43 = 0.26


Area = 0.121 432= 224 in2

Area of segment of circle (2) width chord BC:


2 (

48

2 5.76 5)== 8.26 in

H/D= 8.26

43= 0.19


Area = 0.071 432= 131.3 in2

Area of circle (2) = 432= 1849 in.2

Area available for holes = 1849 (244+ 131.3) = 1473 in.2

Area required for holes = 1634

5holes/in2= 326.8 in2.

Number of holes = 1634

Area = 1473 in.2

23. Tower Height Height of column Hc = (Nact 1)HS +Ht+ Hb +Plate Thickness

No. Of plates = 20

Tray Spacing HS =12 inch

Ht = 5 ft for vapor disengagement

Hb = 2.5ft for liquid hold up

Minimum height of column Htower= (20-1) 1+2.5+5+200.1042

Htower = 26.7 ft

24. Other Mechanical Designs:

Maximum Operating Temperature = 40oC = 313k

Maximum Operating Pressure = 1290 psi = 90.75 kg/cm2

a) Shell Thickness:


Specific Gravity, s = 7.7

Maximum allowable stress, fs = 1300 kg/cm2


Design Pressure, PD = 1.5 Operating Pressure = 1.5 90.75 = 136 kg/cm2

Internal Shell or column Dia, D = 4 ft = 1.2192 m = 1219.2 mm

Here, C = corrosion Allowance = 3 mm

Thickness, ts= PDD

2 fsj PD + C =

1361219.2

2 1300.9 136+ 3 = 78.23 mm

ts= 78.23 mm

b) Tray:

Thickness = 1/8 inch = 0.125 inch


Sieve Tray, hole size = 3/16inch, pitch =1/2 inch, Triangular.

Tray Spacing = 12 inch

c) Support for trays:

A total of 5 inch on diameter for extension of tray ring-type support into the tower.

Material: Carbon steel.

d) Downcomer and Weir:

Material = Carbon steel.

e) Support for Column:

Skirt type, Height = 10 ft , Material = Carbon steel.

f) Heads:

Material = SA-516-70 Carbon steel.

Permissible tensile stress = 1300 kg/cm2

Thickness of Head,

th = PDRcW

2 fsj

=1361219.21.54

2 1300.9

= 109 mm

th = 4.16 inch

Rc = Crown Radius =1219.2 mm

W = Stress intensification factor = 1

4

3 + RcRl

= 1.54 mm

Where, Rl = Knuckle radius = 0.1 Rc

g) Insulation and Coating:

Insulating material: Polyurethane

Insulating width: 4 inch

Coating material: Zinc Sulfate

Coating width: 0.06 inch

Mechanical Design

Vertical section of gas liquid separator

8.7 ft

12.0 ft

44.0 ft

Cross section view of gas liquid separator

8.7 ft

12.0 ft

Heat Exchanger

Heat Exchanger

P & ID

Figure: P & I diagram for centrifugal pump

E-1

P-1

P-2

P-3

feed

Vapor

phase

Liquid phase

AT

AC

0-100

deg

mV

10-50

I-3

P-4 P-2

S-1

S-2

S-3

Xsp

V-1

P-5

S-4

P-6

LIP-7

P-8

light

Figure: P & I diagram for gas liquid separator

V 5 V 4

V 2

V 3

Y strainer

V 1

P

PUMP

TT

TI

AC P/I

Preheated glycol

Heated glycolTo absorber

Dry gas

Sales gas

Figure: P & I diagram for Heat exchanger

Plot Plan

Economic Analysis

14.1 Estimation of Total Capital

Investment and Production Cost

Table 12.1: Equipment cost for Heat Exchanger

Table 12.2: Equipment cost for Pumps

Equipment Identification No. Quantity US $

Centrifugal

Pump P-101 2 9400

Centrifugal

Pump P-100 3 12900

Total = 22300

Table 12.3: Equipment cost for Compressor

Equipment Identification

No. Quantity US $

Reciprocating Compressor K-100 2 35600


No. Quantity US $

Floating head HE E-100 1 80,000

Floating head HE E-102 1 158,000

Total= 238,000

Table 12.4: Equipment cost for Storage Tanks


No. Quantity US $

Storage Tank V-106 1 6300

Table 12.5: Equipment cost for Adsorption Column


No. Quantity US $

Glycol Absorber T-101 4 220,000

DEA Contractor T-100 4 260,000

Total = 480,000

Table 12.6: Equipment cost for Regenerator Column


No. Quantity US $

Glycol Stripper V-103 1 145000

DEA Stripper V-102 1 19000

Total = 164000

Table 12.7: Equipment cost for Stabilizer


No. Quantity US $

Stabilizer V-105 1 7600

Table 12.8: Equipment cost for 2 phase Separators

Total Equipment cost in the year 2002 = $ 958300

Total Equipment cost at present (2014) = $ 958300index value at present

index value at 2002

= $ 958300 443.26

390.4

= $ 1.088 million.


No. Quantity US $

Separator V-100 1 2200



Total = 4500

Table 12.9: Estimation of Capital Investment

Items Percentage of Purchased

Equipment Cost

Cost

Million $

Purchased equipment 100 1.088

Purchased equipment installation 47 0.511

Instrumentation and control 20 0.22

Piping (installed) 68 0.74

Electrical (installed) 11 0.12

Building 18 0.196

Yard improvement 10 0.1088

Service facilities 70 0.762

Land 6 0.0653

Total direct plant cost 350 3.81

Indirect Cost

Engineering and supervision 33 0.36

Construction & expenses 41 0.45

Contractors fee 22 0.24

Contingency 42 0.46

Fixed capital investment

(Indirect + Direct cost)

488 5.32

Ref: Plant Design and Economics by Peter and Timmerhaus; page-251

Working capital investment

(15% of total capital investment)

73.2

0.8

Total Capital Investment 561 6.12

Table 12.10: Manufacturing Expenses

(Working day basis 330day/yr)

Overhead (payroll &

plant), Packing, storage

(50% of opt. labor plus

supervision &

maintenance

0.085

Local taxes (9%of Fixed 0.4788

Cost type Item Cost /yr

$ Million

Major Raw Material

Raw Gas 3.75/1000 scf 7.2

TEG 0.75 $/lb 0.82

Direct

DEAmine 0.6 $/lb 0.02

Cost of Raw materials $ 8.04

Operating labor (10% of RM) 0.804

Supervisory & Clerical

labor

(10% of Opt.

labor) 0.0804

Utilities 0.55

Maintenance & repair

(20% of FC)

1.064

Opt. expenses (10% of

resistance & repair) 0.1064

Laboratory charges

(10% of opt. labor) 0.01064

Total direct manufacturing expenses $ 10.66

capital)

Insurance (0.4% of fixed

capital)

0.0213

Total annual indirect manufacturing expenses $ 0.5851

Total manufacturing expenses $ 11.24

Depreciation (10%

of fixed capital for

machinery &

equipment)

0.532

General Expenses Administrative Cost (20% of Operating labor) 0.1608

Distribution of selling cost ( 10% of total

manufacturing expenses)

1.124

Research & development (5% of total expenses) 0.18

Total annual expenses $ 13.24

Economic Analysis

Percent rate of return

Gas production = 2.92 105 ft3/hr = 2.32 109 scf/year

LPG production = 2356 ton/year

Total sale = $31.34 million

Total income =31.34 13.24 = $ 18.1 million

Net profit after tax = $ 15.4 million (15% Tax)

Total investment cost = $ 6.12 million

Depreciation of fixed capital investment (FCI)

Project life = 20 yr

Salvage value at the end of the project life = 0.15.32= $ 0.532 million

Total depreciation = 0.95.32 = $ 4.788 million

Now depreciation by sinking fund method with 15% interest rate,

= Depreciable FCI (A/P, 15%, 20)

= 4.7880.15 1.1520

1.1520 1

= $ 0.765 million/yr

Cash flow diagram

Payback period calculation

Formula uses

F = P(1 + ),

= (1 + ) 1

(1 + )

Assumed MARR is 15%

1st yr

P0= - $6.12million

So, F1= -$7.04 million

Cost = -$13.24 million

Profit = $15.4 million

So, net cash P1 = -$4.88 million

2nd yr

P1 = -$4.88 million





3rd yr

P2 = -$3.452 million





4th yr

P3= -$1.81 million




So, net cash P4 = $0.08million

Here, net positive cash flow produced, which indicate the payback period.

By linear interpolation, actual payback period is 3.96 yr (3 yr 11month 15 day)

5th yr

P4= -$0.08 million

So, F5= $0.92 million



So, net cashP5 = $ 2.252 million

IRR calculation

In annual worth (AW) method:

Total annual revenue, R= $ (15.4 13.24) = $ 2.16 million

AW= -$6.12 million (A/P, i%, 20) + $ 2.16 million + $ 0.532 million (A/F, i%, 20)

i% (assumption) AW($,million)

20 0.903

40 -0.2897

By linear interpolation, for AW is 0 i% = 35.17%> 15% (MARR)

ERR calculation

Ek( P/F, %,k)(F/P, i%, N) = Rk(F/P, %, N-k)

Where,

% = 15 %

N = 20 yr

Using the above equation, i% = 19.67 %> 15% (MARR)

Appendix

Reference

Ludwig, E.E Applied Process Design for Chemical and Petrochemical Plants, 3

volm,2ndedn, Gulf Pub.

Peters and Timmerhaus, Plant Design and Economics from Chemical Engineers,

4thedn, McGraw Hill (1991)

Phillip C. Wankat, Equilibrium stage separations, prentice hall

A.K.M. Abdul Quader, Design and Building of Process Plant, World University

Service.

Bassel, W.D, Preliminary Chemical Engineering Plant Design, Elsevier Pub, New

York (1976)

Rudd, D.F. and C.C Watson, Strategy of Process engineering, Wiley (1968)

Vilbrandt, F.C and C.E Dryden, Chemical Engineering Plant Design, 4thedn,

McGraw-Hill (1959)

Modern cost Engineering: Methods and Data, compiled and edited by Chemical

Engineering , McGraw-Hill(1979)

Ludwig, E.E Applied Project Engineering and Management, 3rdedn,Gulf Pub.

George T. Austin, Shreves Chemical Process Industries, 5thedn, McGraw-Hill (1954)

RifatM.Dakhil, SataK.Ajjam; General Operating Problems and Their Solutions of

Natural Gas Sweetening Process (Amine System).

http://www.lessmytax.com/cost-inflation-index/

http://allinfoindia.com/eLearning/Income%20tax/current%20index%20of%20capital

%20gain.htm

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