Esterification Reactor Special Design

8/9/2019 Esterification Reactor Special Design

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Cairo UniversityFaculty of EngineeringChemical Engineering Dept.

4th

Year Graduation Project

ESTERIFICATION REACTORSPECIAL DESIGNPoly Ethylene Terephthalate Project PET

Supervised by:Dr Ahmed Seliman& Dr Tarek Mostafa

Submitted by:Mohamed Mohsen Abu-Deif

Sec:3

B.N:33

Date:22-7-2010


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Abstract

Polyethylene Terephthalate PET is one of the most important types of

polyesters which is produced worldwide. PET can be produced via directesterification of Terephthalic acid TPA or Transesterification of Dimethyl

Terephthalate DMT. The main route used now is direct esterification of TPA with

Monoethylene Glycol MEG.

This report will include the detailed design of the Esterification reactor by

direct esterification process. This design includes basic design of the reactor giving its

all dimensions (diameter, height, mixer power and heating media type and its flow

rate). The mechanical design of the reactor also is included.


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1. Introduction

The reactor is the core of any chemical industry as it is the step where reaction

occurs and all the units of the plant are designed either to prepare the feed for the

reaction or to purify the product. So the choice of the reactor affects the whole flowdiagram of the plant and affects also the cost of the plant.

PET production by using TPA and ethylene glycol can be performed by two

different schemes. The first one is monomer production which is called BHET then

polymerization of this monomer to give the polymer PET. While the second scheme is

direct polyesterification of both TPA and MEG to give the polymer. The chosen

process for our project CTIEI process follows the second scheme in which the first

reaction occurs in a CSTR reactor to give the monomer while the polyesterification of

the monomer occurs in two stages prepolymerization and final polymerization.

Modern polyester (1)

Esterification reaction is usually reversible reaction but when the products of

the reaction are removed continuously removed from the reaction medium it can

enhance the reaction in the forward direction. The reaction conditions have to be

adapted with this purpose to give the desired product with desired specifications.

The feed of the reactor is terephthalic acid and ethylene glycol mixtures at 160

C as slurry where the molar ratio of the fresh ethylene glycol to the fresh TPA within

the range 1.05 to 3. After the reaction proceeds the formed monomer BHET and the

unreacted TPA is removed from the bottom of the reactor as slurry with small

amounts of ethylene glycol while the vapour phase consists of water vapour and

excess ethylene glycol which is processed to two distillation columns to recover the

ethylene glycol and to be recycled to the process. (2,3)


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2. Reactor conditions:

The reactor temperature is 260 C and it is somehow constant, and the pressure

is ranged from 1 atm to 2 atm, and in another one from 1 to 3 atm. The selected

pressure for the reactor is 1.5 atm to be within the range of whole literatures. Thereaction can be catalytic or non catalytic. Both are available and industrially

performed. When the reaction is non catalytic it becomes self catalyzed by TPA. This

study proceeds by non catalytic esterification. The effluent from this reactor is

directed to the polymerization reactors which performed in melt phase so usually the

polymerization catalysts and side reactions inhibitors are entered in this reactor but

have no effect on the reaction kinetics or the reactor volume since they are in low

quantities ppm ranges. (2,3)

3. Reaction Kinetics

Firstly the reaction kinetic scheme in which the reaction occurs must be at the

same conditions of the reactor of this study because the rate depends on the process

variables like temperature and pressure and reactant concentrations. The reactions of

the scheme are

)(

)(

2

2

2

1

IIOHBHETMEGMHET

IOHMHETMEGTPA

K

K

++

++

These two reactions are the main reactions but actually monohydroxyethyl

terephthalate MHET is an intermediate in the reaction medium and the reaction is

represented by the overall reaction of the form

OHBHETMEGTPA 222 ++

and the available conversion stands for this overall reaction, so the amount and

concentration of MHET is an unknown.

As stated before these reactions are all reversible but at the reaction conditions 260 Cthe water vapor formed in the reaction is vaporized instantaneously and so the

reaction proceeds in the forward direction according to Le Chatilier principle.

Beside these main reactions there are some side reactions like

)(

)(2

22

213

IVOHDimerMHETTPA

IIIOHDimerMHET K

++

+

The reactions (I-II) are of importance kinetically but the other two reactions

(III-IV) are in equilibrium and can be neglected.


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There is another side reaction in the kinetic form is the formation of diethylene

glycol DEG from the ethylene glycol.

molKJOHDEGMEG RKeq

/32 2 =+

This reaction is important because the DEG content in the final PET product is very

limited especially for bottle grade application. So the reaction must be minimized as

much as possible for that reason stabilizer have to be added in the reaction medium to

work as inhibitors for this reaction and any formed amounts of DEG would vaporized

and directed to the distillation system to be removed.

Through this study the considered reactions would be I and II only and the

related rate equations. Although TPA is solid and the reaction proceeds in the liquidphase and so the diffusion should have a role, it was found that the chemical reaction

is the limiting step in this model since the increase in the homogeneous volume can be

neglected.

The rate equations of the two reactions are below, it can be observed that first

reaction is second order reaction for TPA which acts as the catalyst of this reaction

beside its role as reactant while the order of ethylene glycol is zero as a solvent. For

the second reaction again TPA is the catalyst because its acid properties are more

potent than those of the MHET and it remains permanently in the reaction zone and

also zero order for MEG.

MHET

TPA

CKr

Ckrdis

22

211

=

=

The values of rate constants K1 and K2are calculated at the reaction temperature and

give the values of 24.5 (Kg/mol min) and 8.73 (min-1).

Because of existence of solids in the reaction mixture the solubilities of both reactant

and product must be specified. Here the solubility of TPA in the reactant MEG and

solubility of TPA in the product BHET can be calculated from the equations below

)3831

(

)4877

(

exp374)1

(

exp9062)1

(

TBHET

TMEG

BHETkg

molPAdissolvedT

MEGkg

molPAdissolvedT

=

=

after substituting with reaction temperature 260 C in these equations the value of MEG

is 0.96 and hence it can be assumed that the reaction is liquid phase reaction and this

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phase stands for the reactants phase. In case of solubility of TPA in BHET its value is

0.2 and it means that BHET formed is insoluble in the reaction phase that ensures the

forward direction of the reaction. (4, 5)

4.

Reactor Model EquationsTo be able to determine the reactor volume to achieve the required conversion a

model for the reactor must be constructed. The reactor of this study is CSTR or back

mixing model which is built on the assumption that all properties are constant in the

reactor and the same for the effluent properties of the reactor. According to that

assumption high efficient mixing is required to reach complete mixing. Also because

is the reaction is highly endothermic; efficient heat supply system is needed to

guarantee the same temperature of the reactor.

Equations that describe system are material balance, Energy balance, mixer design

calculations and heat requirements of the reactor. All these equations stand for the

basic design chemical engineering design only.

I.

Material balance equations

According to the reactions mentioned above (I-II) the material balance is

performed with the following known parameters inlet TPA, inlet MEG, overall

conversion of TPA. The other parameters are calculated later. The equations as follow

VmassmixtureTotal

MHETK

massmixtureTotal

TPAKMHETMHET outoutinout

+= 2

2

1

VmassmixtureTotal

MHETKBHETBHET outinout

+= 2

II.

Reactor sizing and Optimizing dimensions:

The variables of the above equations are many but the most important ones are

overall conversion and the volume required for this conversion. Hence this reactor

was simulated via Aspen Plus as stoichiometric reactor in which the conversion must

be specified and then all parameters in the above equations are calculated except

MHETout because this is an intermediate and not found in Aspen library and I cant

add it as user defined components because it results in errors in mass balance

calculations when estimating its properties. So the procedure is as follow;

I. Assume an overall conversion and enter other data required for the simulation

process like reactor conditions inlet flow rates....etc

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II. Run the simulation and take the results from the program like BHETout total

mixture mass and mixture density.

III. So now we have 2 equations in 2 unknowns can be solved simultaneously to

get the volume of the reactor

IV. Repeat the calculations above for several selected conversions and draw the

conversion Vs the reactor volume.

After performing these steps we can obtain the below graph.

According to the required overall conversion the required volume is obtained.

Because we need an overall conversion of 95% the volume of the reactor would be

181.5 m3. This volume is the volume of liquid mixture effluent from the reactor.

According to the assumption of H=2D for most esterification reactors the initial

dimensions of the reactor are D=4.87 m, H=9.74 m.

III. Mixer calculations:

As stated aboveit is required to install high efficient mixing tool to make the

reaction mixture homogeneous. Themixing operation depends on many variables like

type of mixture to be mixed, the existence of heating or cooling and the dimensions of

the equipment where the mixer equipped. Firstly all these variables must be clearly

specified before the mixer design.

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Before the design of mixer, it is important to specify the type of mixer

and its geometry. Multi-stage Pitched blade turbine stirrer of two impellers

both of them are pitched blade type is used here because the ratio of H/D=2

and low viscosity mixtures is employed in the vessel. This type of stirrer can be

equipped with 4 baffles in the tank. About the speed of the stirrer; it can be low

speed to medium speed so 60 rpm is chosen for this type. (6)

The second aspect in mixer design is the power required for the shaft.

An initial estimation of the power required can be calculated via the equation

Power= (0.03-1) KW/M3of the volume of the reactor. By substituting with the

volume of the reactor we get a range of the power from 7 to 181.5 KW, this equation

is valid only for medium mixing in case of reaction with heat transfer conditions. (7)

More accurate equation have to be used to determine the power of the stirrer the

equation is P=NpD5N3 w

Where

Np: the power number which obtained from graphs by knowing Reynolds number

D: diameter of the stirrer which assumed to be 0.5 times tank diameter

N: speed of the of the stirrer in rps

: the reaction mixture density either from literature or simulation

To get Np we must calculate Re firstly

ND2Re = where is the mixture viscosity

After calculating Re a graph for getting the power number but these graphs valid only

for vessels of H/D=1 so this value of Np (1.2 from Jeankoplis) can be assumed to be

valid here because the distance between the two stirrers is equal to the tank diameter.

This calculated power for the single stirrer, to get the power for double stirrer

multiply the resultant power by 1.2 to get the power for the double stage pitched blade

turbine stirrer of the value 122 KW. The stirrer dimensions are illustrated in the

equipment design drawing. (6, 12)

IV. Heat requirements calculations:

To get the heat load required for the reactor, a heat balance around the reactor

must be done; general heat balance equation is,

r

n

joutp

m

iinp

HrVTnCQTnc +=+

== 11

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Where;

m: total number of components in the feed stream

n: total number of components in the effluent stream

Cp: the average heat capacity of component at either inlet temperature or outlet

temperature

Tin: temperature difference between inlet temperature and reference one

Tout: temperature difference between outlet temperature and reference one

Hr:heat of reaction at the reactor temperature

First of all is calculating the heat of reaction at the reactor temperature, this can be

calculated according to Hess law as follow

( )

o

r

C

pV

pL

p

pL

pv

p

r

C

OHBHETMEGTPA

nCn

nCnC

nCn

nCnC

OHBHETMEGTPA

++

+

+

+

+

+ +

2

25

2

260

22

)100260(

)25100()25260(

)18325(

)260183()26025(

22

So we must know Hrowhich also will be calculated from the heat of formations of

the reactants and products as follow

Hro=nprHf

opr-nreactHf

oreact

After the value of Hro is calculated the value of Hrat reaction temperature

also would be calculated and finally substitute in the value of Q or net heat load to the

reactor. There is another easier way to perform energy balance for the reactor; it is

from Aspen plus model of the reactor which was used before to perform the mass

balance. By using this model the net heat load is included in the reactor results afterthe simulation is run. The value of heat load required given by

Aspen is more accurate than the calculated manually value.

V. Internal coil Design:

After the heat load is calculated, it is observed that it is

very high 24415492.4 KJ/hr so using jacket isnt enough to

supply heating required due to its low heat transfer area and

inconvenience for the large diameters. For these reasons coil is

7

Spiral Coils


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used instead of jacket to provide the required heat. There are a lot of configurations

for the coil in the reactor the cheapest one of them is spiral coil. Because it has no

limitation to be used in the reactor, it is chosen for this reactor.(6)

Firstly about the heat transfer media used for heating, there are more than one

alternative. The first one is high pressure steam whose pressure exceeds 15 bar.

Beside HPS there are Dow Therm, Xcel Therm and Paratherm. In the case of this

reactor the pressure is 1.5 bar only so the alternative of HPS is totally rejected because

it would increase the process pressure and also increase the hazard for the process.

For the design procedure of the internal coil, the value of diameter of coil is

assumed to be 3 and the loop diameter is assumed to be 0.8 times the tank diameter

and only one coil is used. This assumptions is valid for the H/D=2 i.e. valid for this

reactor.

Number of loops can be calculated via the equation

coil

vesselloops

D

HN

2=

and hence the length of the coil would be

loopsvesselcoil NDL 8.0=

hence the volume of the coil is calculated via the equation

coilcoilcoil LDV

2

4

=

This volume should be considered in the total volume to be

Rcoiltotal VVV +=

Now new dimensions for the reactor are specified due to the increase in volume.

The area of heat transfer of the coil can be calculated by

coilcoilcoil LDA =

Finally in the equation of

mcoilcoil TUAQ =

there are two unknowns Tmand the overall heat transfer coefficient U and Qcoilis the

same calculated Q before. Tocomplete the design the value of U must be calculated.

There is an equation to determine outside heat transfer coefficient h

4.02.0333.0333.065.022

09.0

=

fbafflesvessel

stirrerptirrero

nD

D

k

CND

k

hD

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Allthe physical constants in the above equation are available either in the simulation

results or in the literatures except for the value off

which can be assumed to equal

one, nbaffle

is the number of baffle equipped in the reactor which has been chosen

before as 4 baffles. This equation is valid for the use of blade turbine stirrer for the

ratio of H/D=2 and baffled vessels i.e. it is valid to be used in this case. In case of the

low thickness of the tube, the conduction resistance of the tube can be neglected. For

the case of choosing very efficient heating media in the coils its resistance also can be

neglected. Thats why in this case the calculated h can be assumed as U the overall

heat transfer coefficient.

Returning to the equation of heat transfer rate mentioned above the only unknown

now is logarithmic mean temperature difference Tmso it can be calculated easily. To

find the conditions of the heating medium like flow rate inlet and outlet

temperature, the following must be done.

=

out

in

outin

m

T

TLn

TTT where Rinoilin TTT = and Roilout TTT out =

To get the reasonable Tin of the heating medium, trial and error should be used. The

first trial is assuming the temperature approach of 15 C; it means that outlet

temperature of the heating medium is 275 C. With substituting in the last equation the

value of inlet temperature of the heating medium is calculated as 472 C. It is found

that this temperature is very high to be the inlet temperature of a heating medium so

another trial is required.

The second trial is in reverse way, the heating medium alternatives (Dow

Therm vapour phase, Xcel Therm vapour phase and Xcel Therm liquid phase) are

compared together by their characteristics and properties and it is found that the liquid

phase Xcel Therm is better than the others. Beside the phase of the media there is

another point of comparison is the flow rate required for the heating which is

calculated later. Hence the inlet temperature is 400 C and then substitute again to get

the outlet temperature of the oil is 293 C. (Xceltherm site and Dow Therm file)

Finally the flow rate of the heating medium is calculated from the equation

)( outinp

coil

coil

TTC

QF

=

9


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It results the value of 53 ton/hr; it is observed that this value id=s very high but

reasonable with respect to the volume of the reactor and inlet mass flow rate and heat

load of the reactor. (8, 9, 10)

5.

Mechanical DesignI.

Material of construction

It was noticed that the reactor will be subjected to slurry with acidic medium at high

temperature of 260 C, so the material used must withstand these conditions and

Stainless steel 316 L was found to be suitable in such a reactor. Stainless steel 316L

contains 18% Cr + 8% Ni +0.03% C + 3% Mo. Carbon steel is also reasonable but

less used. (11)

II.

Calculations:

The reactor is a tall vessel under internal pressure and it will be designed at coastal

area. Design according to ASME code.

(1)Design of thickness of the shell (2)Design of dished head and bottom

Pd= 1.1 Pmax = 23.94 psi

Rish= 98.425

all= 18000 psi

Assume: C = 1/16 ; E = 1

tsh= 4/16

Assume r/R=0.1 so K=1.5

R=Dish

th=tb=3/16

(3) Design of neck (4) Gasket deign

existing Necks

manhole+2 other openings each one is 50 cm

diameter

Di neck=19.68 assumed

tneck=2/16

Do neck=19.84

Type steel

M=2.75, y=3700

Assume Dig=Dineck=19.68

Dog=19.75

b=0.26/8


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(5) Anchor bolts Design (6) Flange Design

Dmg=19.81003937

all bolts=12000 psi

c=0.0625

dib=1

Q'b=7231.261721bf

Q"b=14391.8609 bf

n


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6.

Cost estimation of the reactor

There are a lot of factors that affect the cost of the reactors. As being CSTR reactor it

is required to got vessel cost individually, the cost of mixer and the shaft and the

motor power and finally the cost of heating system of the reactor here it is the coil

Pw=0.27

In case of operation:

. = + -

. = - +

= 0

=.

. = -

.= +

In case of shut down:

w=

= 0

=.

. = -

.= +

All the checks are verified and the shell is safe

ith t=4/16

(10) Design of skirt support (11) Design of Bearing plate

Assume D i sk=0.98 Di sh

Do sk=1.02 Do sh

L sk=2.5 m

tsk=(Do-Di)/2+c

performing all checks for tall vessel as

mentioned above

All checks is verified and the skirt is safe

with the selected dimensions

Assume

Repeat the checks for the wind will result in

verifying the checks and the plate is safe with

assumed dimensions

(12) Design of insulation (13) Design of shaft

Type of insulation glasswood

Thickness of 1 is enough

The motor shaft must be supported because of

the large height in the vessel.

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system. That is all about fixed capital cost of the reactor. About the

variable cost of the of the reactor it is required to specify the power of the

mixer and the heating medium flow rate. Beside these accurate method of

the cost estimation, there are shortcut methods like estimation of the

reactor cost by its dimensions only and multiply this cost by a factor to

get the estimation of the mixer and the heating system. Attached below

the cost estimation of the process including the reactor. (13, 14)

ECONOMIC STUDY

1. Estimation Of Purchased Cost Of Equipment (PCE):

Table (1): Purchased Cost of Equipment for TPA production (Eastman)

Eastman TPA

EquipmentCost At2007 ($)

Total Costs($)

CompressorsCompressor1 181,500

Compressor2 197,700

Compressor3 128,400

Compressor4 383,000

Total Compressors Cost 890,600

ExpandersExpander1 130,000

Expander2 98,000

Total Expander Cost 228000

Flash SeparatorsOver Flow Flash Drum 68,200

Vacuum Flash Drum 80,100

Total Flash Cost 148,300

WRCCost Of Top Trays 87,913.8

Cost Of Bottom Trays 36,414

Top Shell 124,100

Bottom Shell 32,700

Total WRC Cost 281127.8

Heat ExchangersInter-Stage Cooler1 1,405,300

Inter-Stage Cooler2 1,303,000

WRC Condenser 265,600

Flash Condenser 2,437,400

Rankine Condenser 328,000Solvent Vaporizer 82,800

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Pre Expander Gas Heater 353,900

After Crystallizer Condenser 202,500

After Over Flow Flash Drum Condenser 511,700

Total Heat Exchangers Cost 6,890,200

Dryer 479,700

Lot Bin 576,300

TanksFiltrate Tank 558900

Reflux Tank 470800

Total Tanks Cost 1,029,700

Table (1 contd.): Purchased Cost of Equipment for TPA production (Eastman)

Table (2): Purchased Cost of Equipment for Intermediate PET production(CTIEI)

Eastman TPA

EquipmentCost At 2007

($)

Total Costs

($)Static Mixer 8,500

Centrifuge 203,500

Rotary Vacuum FilterRotary Vacuum Filter 346,500

RVF Vacuum System 13,800

Total RVF Cost 360,300

Vacuum Flash Drum Vacuum System 10,300

Crystallizer 1,184,000

Post Oxidizer

Shell 306,300Impeller 134,300

Cost of 1 stage 440,600

Total Post Oxidizer Cost 881,200

Bubble Column OxidizerTitanium lining 4,778,100

Carbon steel shell 5,032,300

Total Bubble Column Cost 9,810,400

Total Fixed Cost 2007 USD$ 22,982,127.82007 Cost Index 525.4

2010 Cost Index 555.2

Total PCE For TPA At 2010 24,285,643.99

CTIEI


($)

Total Costs

($)First Distillation Column

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Table (2 contd.): Purchased Cost of Equipment for Intermediate PETproduction (CTIEI)

Table (3): Purchased Cost of Equipment for Product PET production (Buhler)Buhler


($)Total Costs

($)

Total purchasing cost Buhler 2007 6,431,103



Total purchasing cost Buhler 2010 6,795,866.741

Total purchased cost for project 35,612,088.56

First Column 38,600

First Column Condenser 47,300

First Column Trays 5,357

Total First Distillation Cost

91,257

Second Distillation Column

Second Column 69,800

Second Column Condenser 35,400

Second Column Reboiler 91,800

Second Column Trays 3,689

Packing 22,406

Total Second Distillation Cost

223,095

Vacuum Unit

20,700

Pre-Polymerization ReactorPre-Polymerizer 786,500

Agitator 247,000

Total Pre-Polymerizer Cost

1,033,500

Esterification Reactor

Esterifier 1,117,700

Agitator 45,200

Total Esterifier Cost

1,162,900

Mixer

Agitator 36,200

Vessel 43,200Total Mixer Cost

79,400

Disk Ring Reactor

1,676,550

Total Purchasing Cost CTIEI 2007 ($) 4,287,402



Total Purchasing Cost CTIEI 2010

4,530,577.827

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2.

Estimation of Total Capital Investment:

Table (4): Fixed and Working Capital Investment

Fixed Capital Investment

Item Fluid - Solid factor Factor*PCEPurchased Cost of Equipment X 35,004,148.99

installation 0.45 x 15,751,867.05

Piping 0.45 x 15,751,867.05

instrumentation 0.15 x 5,250,622.349

Electrical 0.1 x 3,500,414.899

Building 0.1 x 3,500,414.899

Utilities 0.375 x 13,126,555.87

Storage 0.2 x 7,000,829.798

Site development 0.05 x 1,750,207.45

Ancillary buildings 0.225 x 7,875,933.523

PPC 3.1 x 1,0851,2861.9

Design and Engineering 0.25 PPC 27,128,215.47

Contractor's fees 0.05 PPC 5,425,643.093

Contingency 0.1 PPC 1,0851,286.19

Total 43,405,144.75

FCI 151,918,006.6

Working Capital InvestmentWCI 0.2 FCI 22,787,700.99

Total Capital InvestmentTCI FCI + WCI 174,705,707.6

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3.

Estimation of Net Profit/Year and Payback Time:

Table (5): Cost of Raw Materials and Products

Sales And Raw Materials Cost

Material Price Per Ton CostPX 1200 301,468,800

TPA 925 92,500,000

PET 1500 525,000,000

AcAc 530 288

MEG 550 70,400,000

Table (6): Price and Cost of Utilities

UtilitiesType Price ($) Price Unit Amount Cost

Steam 0.355 Per Ton Product 124,250 6,833,750Cooling 0.251 Per Ton Product 87,850 8,785,000

Electricity 96kWh Per Ton

Product 33,600,000 6,048,000

Fuel 500 m3/hr 4,000,000 1,400,000

Total Cost ($) 12,581,864

Table (7): Calculation of Net Profit/year and Payback time

Manufacturing Cost

Direct Cost Calculation Equation Value ($)A- Variable Cost

Raw Materials 371,869,088.00

Miscellaneous 0.1 Maintenance 1,519,180.07

Utilities 12,581,863.64

Shipping And Packing $ 14per Ton 4,900,000.00

Total Variable Cost 390,870,131.70

B- Fixed Cost

Labor $ 1000Per Month * 150 Employee 1,800,000.00

Supervision 0.2 Labor 360,000.00

Plant Overhead 0.5 Labor 900,000.00

Depreciation 0.15 FCI 22,787,700.99

Interest 0.01 FCI 1,519,180.07

Insurance 0.01 FCI 1,519,180.07

Rent 0.01 FCI 1,519,180.07

Royalties 0.01 FCI 1,519,180.07

Maintenance 0.1 FCI 15,191,800.66

Total Fixed Cost 47,116,221.92

Direct Production Cost Variable Cost + Fixed Cost 437,986,353.62

Table (7 contd.): Calculation of Net Profit/year and Payback time

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Manufacturing Cost

Indirect Cost Calculation Equation Value

Indirect Production Cost 0.25 Direct Production Cost 109,496,588.41

Manufacturing CostDirect Production Cost +Indirect Production Cost 547,482,942.03

SalesSelling: 100,000 Ton TPA 350,000 Ton PET

617,500,000.00

Gross Profit Sales - Manufacturing Cost 70,017,057.97

Net Profit Per Year At Taxes 20% of Gross Profit 56,013,646.38

Payback Time (years) TCI/Average Profit Per Year 3.12

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References:

1. Modern Polyesters:Chemistry and Technology of Polyesters and Copolyesters.

John Scheirs and Timothy E. Long, John Wiley & Sons Ltd.

2.

Organic chemical process industry 9/91 (reformatted 1/95)

3. United states patents, patent no 6096838, methods and apparatus for

continuous polycondensation Assignee; Hitachi limited, Japan.

4. Differential kinetic model of the polyesterification of Terephthalic acid with

ethylene glycol without catalyst, M.P.Vladimirova, V.V.Kiselev &

V.A.Malykh, and A.S.Chegolya.

5. Identification of kinetics of Direct Esterification Reactions for PET Synthesis

Based on a Genetic Algorithm, Department of Chemical Engineering,

Hanyang University.

6.

Stirring: Theory and Practice, Marko Zlokarnik, Wiley VCH.

7. Coulson and Richardson Chemical Engineering Volume 6, R.K.Sinnott.

8.

Chemical reactor design and control, William L. Luyben,Lehigh University,Wiley 2007.

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Esterification Reactor Special Design

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