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Cairo University Faculty of Engineering Chemical Engineering Dept. 4th Year Graduation Project

ESTERIFICATION REACTOR SPECIAL DESIGN Poly 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 direct esterification 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.

I|

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 flow diagram 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. The reaction 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

3

K1 TPA + MEG → MHET + H 2 O

(I )

K2 MHET + MEG → BHET + H 2 O

( II )

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

TPA + 2MEG → BHET + 2 H 2 O 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 C the 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 K3 2MHET → Dimer1 + H 2 O

( III )

TPA + MHET ← → Dimer2 + H 2 O

( IV )

The reactions (I-II) are of importance kinetically but the other two reactions (III-IV) are in equilibrium and can be neglected.

2

There is another side reaction in the kinetic form is the formation of diethylene glycol DEG from the ethylene glycol. eq 2 MEG ←  → DEG + H 2 O

∆Η R = −3KJ / mol

K

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 liquid phase 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. 2 r1 = k1CTPA dis

r2 = − K 2 C MHET The values of rate constants K1 and K2 are 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 ( dissolvedTPA mol α MEG ( ) = 9062 exp 1kg MEG

α BHET (

( dissolvedTPA mol ) = 374 exp 1kg BHET

−4877 ) T

−3831 ) T

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

Reactor Model Equations

4.

To 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 3 (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

MHETout

2       TPAout MHETout   − K 2   × ρ ×V = MHETin + K1    Total mixture mass  Total mixture mass     

  MHETout  × ρ × V BHETout = BHETin + K 2   Total mixture mass 

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 can’t 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. 250 230

Reactor Volume cum

210 190 170 150 130 110 90

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70 50 0.4

0.5

0.6

0.7

0.8

0.9

1

Overall Conversion

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 above it is required to install high efficient mixing tool to make the

reaction mixture homogeneous. The mixing 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/M3 of 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=NpD5N3ρ

w

Where Np: the power number which obtained from graphs by knowing Reynolds number D: diameter of the stirrer which assumed to 3be 0.5 times tank diameter N: speed of the of the stirrer in rps ρ: the reaction mixture density either from literature or simulation To get Np we must calculate Re firstly

Re =

D 2 Nρ

µ

where µ 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, m

n

i =1

j =1

∑ nc p ∆Tin + Q = ∑ nC p ∆Tout + rV∆H r

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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 TPA +

2 MEG

↓

↓

nC p (25 − 260) ↓ TPA +

C 260   →

 nC pv (183 − 260) +     nλ + nC (25 − 183)  pL   ↓

2 MEG

C 25 → 

BHET +

2 H 2O

(nC

 nC pL (100 − 25) +     nλ + nC (260 − 100)  pV   ↑

p

(260 − 25) ) ↑

3

BHET +

2 H 2O

∆Η r

∆Η or

So we must know ∆Hro which also will be calculated from the heat of formations of the reactants and products as follow

∆Hro=∑npr∆Hfopr-∑nreact∆Hforeact After the value of ∆Hro is calculated the value of ∆Hr at 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 after the 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 isn’t enough to supply heating required due to its low heat transfer area and inconvenience for the large diameters. For these reasons coil is

Spiral Coils 72

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 N loops =

H vessel 2 Dcoil

and hence the length of the coil would be Lcoil = 0.8 Dvesselπ N loops

3

hence the volume of the coil is calculated via the equation

Vcoil =

π 4

2 Dcoil Lcoil

This volume should be considered in the total volume to be

Vtotal = Vcoil + VR 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

Acoil = π Dcoil Lcoil Finally in the equation of

Qcoil = UAcoil ∆Tm there are two unknowns ∆Tm and the overall heat transfer coefficient U and Qcoil is the same calculated Q before. To complete the design the value of U must be calculated. There is an equation to determine outside heat transfer coefficient h  D 2 Nρ  hDo  = 0.09 tirrer k µ  

0.65

 Cpµ     k 

0.333

 Dstirrer   Dvessel

  

0.333

 2  n  baffles

   

0.2

 µ  µ  f

   

0.4

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All the physical constants in the above equation are available either in the simulation results or in the literatures except for the value of

µ which can be assumed to equal µf

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. That’s 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 ∆Tm so 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.

∆Tm =

∆Tin − ∆Tout where ∆Tin = Toil in − TR and ∆Tout = Toilout − TR  ∆Tin   Ln  3 ∆ T  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

Fcoil =

Qcoil C p (Tin − Tout )

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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 Design

I.

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 3

Pd = 1.1 Pmax = 23.94 psi Rish = 98.425’’ σall = 18000 psi Assume:

C = 1/16 ‘’ ; E = 1

Assume r/R=0.1 so K=1.5 R=Dish th =tb =3/16 ‘’

tsh = 4/16 (3) Design of neck

(4) Gasket deign

existing Necks

Type steel

manhole+2 other openings each one is 50 cm

M=2.75, y=3700

diameter

Assume Dig=Dineck=19.68” Di neck=19.68” assumed

Dog=19.75”

tneck=2/16 “

b=0.26/8”