Apparatus design project on heat exchanger

Apparatus Design Project on Heat Exchanger and Distillation column Design By Abnet Mengesha 2017
i
Acknowledgment
I wish to acknowledge and thank all my supporters in journey of my project work. In particular, I
would like to express my deepest gratitude and appreciation to Dr. Solomon Bogale and Mr.
Kalid our project advisor and lecturers for their limitless support and giving advice throughout
the semester.
I am very grateful to my class mates for their supports, giving necessary materials and data,
sharing ideas during my project work.
In closing, I extend my gratitude to my friend Fitala Bula for his willingness to support me by
giving idea and his laptop for my project work.

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Executive summary
Most of the processes in chemical industry are involved in purifying components. As a
consequence, a large part of the energy use in many industrial sectors can be attributed to
separation processes.
Distillation is the dominant separation technology in chemical industries despite its huge energy
consumption. Distillation consumes about 3% of the total energy consumed globally.
Since many separation tasks need to continue with this technology, methods to determine the
minimal energy used in a given distillation task have become important. For separations of a
multicomponent mixture, one way to reduce the energy requirements is using thermally coupled
distillation columns instead of the conventional direct sequence. These new methods permit
energy savings more than 30% in comparison with conventional sequence distillation columns.
Moreover, the recent rise in energy prices and demands further emphasizes the relevance of this
problem.
Apparatus design is one of the course that given to the fourth year chemical engineering students
at Addis Ababa University AAIT. Here we are given a design project that will be done
throughout the semester. Thus, the task is given to me as to design the distillation column and its
corresponding preheater for acetone water mixture be separated as top and bottom products.
This part of the report aimed to explain the work that have been done by me as the design project
throughout the full semester. The report divided into two main parts.
The first part is all about the design of preheater to facilitate the coming separation process of
acetone water mixture. Here, heat source is selected, thermal design is done, pressure drop both
in tube and shell is determine.
The second part is concerned with the design of distillation column for acetone water mixture to
be separated. In this part, column height, diameter, pressure drop per tray and column internal
design is completed.

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Abbreviation and Acronym
The most important nomenclature used in this report can be summarized in:
w/w mass by mass fraction
Q heat transfer rate
M mass flow rate
T hot fluid temperature
t cold fluid temperature
Acp area per pass
Ut tube side velocity
Us shell side velocity
Db tube bundle diameter
Ds shell inside diameter
Lb baffle spacing
∆Pt tube side pressure drop
∆Ps shell side pressure drop
V Vapor flow rate
VT Vapor flow rate in the top
L Liquid flow rate
D Distillation product
B Bottom product
q Liquid fraction
z Mole fraction in feed
α Relative volatility
NC number of components

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TABLE OF CONTENTS
Acknowledgment ............................................................................................................................. i
Executive summary.........................................................................................................................ii
PART I..........................................................................................................................................vii
HEAT EXCHANGER DESIGN..................................................................................................... 1
1Introduction................................................................................................................................... 2
1.1Problem Statement ................................................................................................................. 3
1.2 Objectives.............................................................................................................................. 3
1.2.1 General Objectives ......................................................................................................... 3
1.2.2 Specific Objectives......................................................................................................... 3
1.3 Scope of the Project............................................................................................................... 3
1.4 General Assumptions, Notes and Facts................................................................................. 4
2 Heat Source Selection.................................................................................................................. 5
3 Thermal Design............................................................................................................................ 6
3.1 Specifications and Assumptions............................................................................................ 6
3.1.1Energy Balance................................................................................................................ 8
3.2 Collection of Physical Properties .......................................................................................... 9
3.2.2 Calculation of Mean Temperature & Fluid Properties at mean temp........................... 10
3.3 Overall Coefficient.............................................................................................................. 11
3.4 Heat Exchanger Type and Dimensions ............................................................................... 11
3.4.1 Log Mean Temperature and True Mean Temperature ................................................. 12
3.5 Heat Transfer Area (A0) ...................................................................................................... 12
3.6 Exchanger Layout and Tube Size........................................................................................ 13
3.7 Numbers of Tube................................................................................................................. 13

v
3.8 Tube Bundle Diameter and Shell Diameter (Db.)................................................................ 14
3.9 Tube-Side Heat Transfer Coefficient (hi)............................................................................ 14
3.10 Shell Side Heat Transfer Coefficient ................................................................................ 15
3.11 Overall Heat Transfer Coefficient..................................................................................... 16
3.12 Pressure Drop .................................................................................................................... 16
3.12.1 Tube Side.................................................................................................................... 16
3.12.2Shell Side Pressure Drop (∆Ps).................................................................................... 16
4 Summary of Notes Taken from the Literature........................................................................... 19
5 References.................................................................................................................................. 22
PART II......................................................................................................................................... 23
DISTILLATION COLUMN DESIGN......................................................................................... 23
5 Introduction................................................................................................................................ 24
5.1 Problem Statement .............................................................................................................. 25
5.2 Mass Balance and Determination of No of Theoretical Stage ............................................ 28
5.3 Estimation of Physical Properties ....................................................................................... 31
5.4 Plat Spacing......................................................................................................................... 32
5.5 Column Diameter Estimation.............................................................................................. 32
5.6 Selection of Liquid-Flow Arrangement .............................................................................. 35
5.7 Make Provisional Tray Layout............................................................................................ 35
5.8 Check the Weeping Rate..................................................................................................... 36
5.9 Check Plate Pressure Drop.................................................................................................. 38
5. 10 Down-Comer Backup Liquid and Down-Comer Residence Time .................................. 39
5.10.1Down-comer design [back-up].................................................................................... 39
5.11 Plate Layout....................................................................................................................... 40

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5.12 Flooding and Entrainment Checking................................................................................. 41
6 Conclusion ................................................................................................................................. 42
7 Reference ................................................................................................................................... 43

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List of Tables
Page
Table 1: Data for first heat exchanger (heat source top product) ................................................... 6
Table 2: Data for the second heat exchanger (heat source bottom product)................................... 7
Table 3: Some assumptions and facts ............................................................................................. 7
Table 4: At inlet temperature.......................................................................................................... 9
Table 5: Physical properties at mean temperature ........................................................................ 10
Table 6: Feed and product composition........................................................................................ 28
Table 7: Summary of material balance around the distillation column........................................ 31

viii
List of Figures
Page
Figure 1: Provisional area and dimension..................................................................................... 35
Figure 2: Relation between down comer area and weir length..................................................... 35
Figure 3: Weep-point correlation.................................................................................................. 38

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PART I
HEAT EXCHANGER DESIGN
PROJECT ON ACETONE WATER MIXTURE
PREHEATER DESIGN
BY ABNET MENGESHA DUBE
ADDIS ABABA ETHIOPIA, JUNE, 2017

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1Introduction
Heat exchangers are devices that facilitate the exchange of heat between two fluids that are at
different temperatures while keeping them from mixing with each other. Heat exchangers are
commonly used in practice in a wide range of applications, from heating and air-conditioning
systems in a household, to chemical processing and power production in large plants.
Heat transfer in a heat exchanger usually involves convection in each fluid and conduction
through the wall separating the two fluids. In the analysis of heat exchangers, it is convenient to
work with an overall heat transfer coefficient U that accounts for the contribution of all these
effects on heat transfer. The rate of heat transfer between the two fluids at a location in a heat
exchanger depends on the magnitude of the temperature difference at that location, which varies
along the heat exchanger.
Perhaps the most common type of heat exchanger in industrial applications is the shell-and-tube
heat exchanger. Shell-and-tube heat exchangers contain a large number of tubes (sometimes
several hundred) packed in a shell with their axes parallel to that of the shell. Heat transfer
takes place as one fluid flows inside the tubes while the other fluid flows outside the tubes
through the shell. Baffles are commonly placed in the shell to force the shell-side fluid to flow
across the shell to enhance heat transfer and to maintain uniform spacing between the tubes.
In this project work, preheater for acetone water mixture is needed to be designed. The preheater
aimed to heat the mixture form its initial condition (20 0
C and around 1 atmosphere) to its final
state to saturated liquid at bubble point temperature. In the process of designing, several tasks are
accomplished. The major and the first task is selection of heat source and determination of its
inlet and outlet temperature, the second task is to collect different thermos-physical properties
and performing some calculations. And the last task is to determine the tube and shell layout,
thermal design part is done and pressure drop for both tube and shell side fluid is calculated.
Finally, the result is discussed in spread sheet form and the optimum design also discussed in the
same form to compare the design task and the optimum one.

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1.1Problem Statement
1500kg/hr. of 10/90 (w/w) % acetone water mixture will be heated from 20 0
C to t2, to prepare
the mixture for the distillation process that produces top product of 90% (w/w) acetone and the
rest water, and the bottom products of (1/99) % (mole/mole) acetone/water. Design heat
exchanger for this service.
Note: Here the designer must, select the heat source, determine feed outlet temperature,
determine heat source’s inlet and outlet temperature, decide types of heat exchanger, heat duty,
tube area, tubes number, tube layout, baffle dimension, shell dimension and material types.
1.2 Objectives
1.2.1 General Objectives
The general objectives of this project are:
 To help students to use the knowledge and skill they gained from 1st
year to now.
 To help student to search their real potential through the design project.
1.2.2 Specific Objectives
The specific objectives of this project are:
 To design economical and efficient preheater for acetone water mixture
 To calculate economical area
 Optimization between pressure drop and heat transfer coefficient through thermal design
1.3 Scope of the Project
This project is a preliminary design of preheater, by the fourth year chemical engineering
students of Addis Ababa University technology institute as their semester project. Since the
project is not detailed one it is not advisable to use it as a guide or other material at your detailed
design.

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1.4General Assumptions, Notes and Facts
 Water and acetone is missile but negatively deviated from rault’s law, thus it is not ideal
solution.
 Water and acetone has 40 boiling point differences at their normal boiling point, thus it
required les amount of energy to separate them at this state.
 Let the outlet temperature of the mixture is the bubble point temperature of the mixture,
at atmospheric pressure 830
C, this is needed to facilitate the coming separation process.
 The cold and the hot streams generally, introduced at atmospheric pressure to the
exchanger, at some point the hot stream pressure may change. At this atmospheric
pressure the saturation temperature of water is 100 0
C
 The mixture bubble and dew point temperature is 83 and 95 0
C respectively. The bubble
point temperature has selected as outlet temperature of heat exchanger.
 Since its suitability and availability shell and tube heat exchanger has selected. Flow
arrangement and other issues are addressed based on the result we will obtain through the
calculations.
Shell and tube heat exchanger have selected because of the following advantages. The
advantages of this type exchanger are:
1. The configuration gives a large surface area in a small volume.
2. Good mechanical layout: a good shape for pressure operation.
3. Uses well-established fabrication techniques.
4. Can be constructed from a wide range of materials.
5. Easily cleaned.
6. Well-established design procedures.
NB: It is generally the case that a feed (at the column pressure) should enter the distillation
column as a saturated liquid or at a temperature between its bubble point and dew point
temperatures. A superheated feed or a sub cooled feed will introduce thermodynamic
inefficiencies, as it will need to be cooled or heated, respectively, to saturation conditions to

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participate in the separation processes within the column. Thus I have selected the bubble point
temperature as outlet temperature of the feed fluid from heat exchanger and the inlet temperature
of distillation column.
2Heat Source Selection
Four heat sources are selected as an option to preheat acetone water mixture before it introduced
into the distillation column. The process is continuous and some pumping cost is may added. The
sources are:
1. The top product vapor, which is at atmospheric pressure and 58.20
C. (the feed outlet
temperature is greater than this the source can change the feed temperature in little
amount)
2. The bottom product, which is at1.2533bar pressure and 1060
C (optimum)
3. Condensate (saturated liquid water) from re-boiling section, which is about 180-200
0
C(it’s better to use this source as feed water to boiler section)
4. Raw steam (superheated), since it has high quality, it is very expansive and to use this
source as preheating agent is uneconomical.
Among these four heat source the bottom product have selected as major heat source, since it has
enough energy that is needed for the process. Even if the bottom product is my major heat
source, latent heat from the top product also taken to rise the feed’s temperature to some extent.
The other sources are not selected based on steam economy and heat exchanger size.

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3Thermal Design
3.1 Specifications and Assumptions
Table 1: Data for first heat exchanger (heat source top product)
Properties Hot
fluid
Hot
acetone
Hot water Cold fluid Cold
Acetone
Cold Water
Mass fraction 1 1 0.1 0.9
Inlet temperature
0
C
58.2 58.2 58.2 20 20 20
Outlet temp 0
C 58.2 53 53 53
Mass flow rate
kg/hr.
279.9 1500 150 1350
Vapor mole
fraction
0.8856 0.1143
Latent heat of
vaporization
kJ/kg
707 501.6 2300

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Table 2: Data for the second heat exchanger (heat source bottom product)
Properties Hot fluid Cold water Cold acetone Cold fluid
Mass fraction 1 0.9 0.1 1
Inlet temp 0
C 106=T1 53=t1 53 53
Outlet temp
0
C
87(iterative result)=T2 83=t2 83 83
Mass flow
rate kg/hr.
1353.6=B (material balance on
distillation column)
1350 150 1500
Table 3: Some assumptions and facts
Given data/assumptions Hot fluid Cold fluid
Fouling factor (m20
C/w) 0.0003-0.0002 0.0003-0.0002
Inlet pressure (bar) 1.2133 1.01325
Allowable Pressure drop (bar) ≤ 0.35 ≤0.35
Mass fraction (w/w) %, for cold mixture Water Water/acetone=90/10
Mole fraction 0.01 acetone 0.0333 acetone

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3.1.1Energy Balance
QC= M c *∆Tc *Cpc=84.78754kw (this amount of energy is required to rise feed temperature
from 200
C to 830
C).
To use the bottom product only as heat source, cause the temperature cross to be occurred. Thus
for this reason and to reduce condenser load and cooling water requirement, I have used two heat
exchanger sequential, one that use the top product as heat source and the other that use bottom
product. Here I have performed design for the latter heat exchanger only and for the first heat
exchanger a little calculation is done to find T2.
Qh1 = Mh *hfg = 707kJ/kg * 279.9kg/hr. = 55kw
Qc1=Qh= Mc *Cpc1*∆Tc1
Qh1/ Mc = 132kJ/kg = Cpc1*∆Tc1
As first trial take the outlet temperature of the cold fluid as equal to the outlet temperature of the
hot fluid (580
C), CPc1 = 3.995 kJ/kg.k. at mean temperature (39). Thus,
132kJ/kg = (3.995kJ/kg. k) *(t2-20)
t2 = 530
C and stream mean temperature is (53 +20)/2 = 36.5 and at this temperature CPc1 =
3.98737 kJ/kg k. Thus,
132kJ/kg = 3.98737*(t2-20)
t2 = 53.10
C
There is no significant change in the specific heat at this mean temperature from the value used,
so take the cold stream outlet temperature to be 53.10
C, say 530
C. (this temperature is the outlet
temperature of the first preheater and the inlet temperature of the second preheater).
Qc2 = Qc – Qc1= 84.7875-55
Qh2 = 29.7875kw = B* Cph2*∆Th2
(29.7875kJ/s)/ (0.376kg/s) = 79.22kJ/kg = Cph2*∆Th2
As first trial take the mean temperature of the hot fluid as equal to the outlet temperature of the
cold fluid (830
C) at this temperature CPc1 = 4.1973 kJ/kg.k. Thus,

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79.22kJ/kg = (4.1973kJ/kg. k) *(106-T2)
T2 = 87 0
C and stream mean temperature is (87+106)/2 = 97 and at this temperature CPh2 = 4.214
kJ/kg k. Thus, 79.22kJ/kg = (4.214kJ/kg. k) *(106-T2), T2 = 87.2 0
C
There is no significant change in the specific heat at this mean temperature from the value used,
so take the hot stream outlet temperature to be 87.2 0
C, say 87 0
C.
Where
Mc=mass flow rate of the cold fluid
Mh= mass flow rate of the hot fluid
∆T, inlet and outlet temperature difference for both cold and hot fluid
Cp, specific heat for both fluids
3.2Collection of Physical Properties
Table 4: At inlet temperature
Properties Hot fluid @
106 0
C
Cold fluid @
53 0
C
Acetone Cold water
Density kg/m3
954.708 957.9134 756.16 987.2
Viscosity kg/m.hr 0.9634 1.9 1.855 1.905
Specific heat kJ/kg.k 4.223 3.99 2.2642 4.182
Thermal conductivity w/m.k 0.6805 0.59454 0.1494 0.644

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3.2.2Calculation of Mean Temperature & Fluid Properties at mean temp
Table 5: Physical properties at mean temperature
Fluid Inlet temperature 0
C Outlet temperature 0
C Mean temp/film0
C
Hot 106 87 97
Cold 53 83 68
Properties Hot fluid (@
97)
Cold fluid Water (@ 68) Acetone(@ 68)
Viscosity (kg/m.hr.) 0.9612 1.607 1.62 1.5
Density (kg/m3
) 960 947.3 978.66 735.4
Thermal conductivity
w/m.k
0.678 0.61 0.6614 0.147
Specific heat capacity
(kJ/kg.k)
4.214 4.004 4.189 2.3382
Different formulas to calculate the properties of acetone water mixture @ 680
C
Mixture density (p)
1/ ρ=x1/ ρ1+x2/ ρ2
Where ρ=mixture density
ρ1=component 1density
ρ2= component 2 density
x1&x2 mass fraction of component 1&2 respectively

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Viscosity of mixture (µm)
1/µm=x1/ µ1+x2/ µ2
Where µm, µ1 & µ2 viscosity of mixture, component 1&2 respectively
Thermal conductivity of mixture (Km)
Km =x1k1+x2k2
Where
K1 &K2, thermal conductivity of component 1&2
Specific heat capacity of mixture (Cpm)
Cpm=X1Cp1+x2Cp2
Where
Cp1 &Cp2, specific heat capacity of component 1&2
3.3 Overall Coefficient
From the literature the overall heat transfer coefficient for this type of fluid (i.e. bottom product
almost water and also feed fluid dominated by water) is ranged between 800-1500 w/m20
C
Let start the calculation at the middle of the two values (1000w/m20
C)
3.4 Heat Exchanger Type and Dimensions
An even number of tube passes is usually the preferred arrangement, as these positions the inlet
and outlet nozzles at the same end of the exchanger, which simplifies the pipe work. Let start
with one shell pass and 4 tube passes. And counter current flow arrangement. Exchanger with an
internal rotating head is more versatile than fixed and U-tube exchanger. It is suitable for
 High temperature difference between shell side and tube side fluids (specially above 80
0
C)
 As the tubes can be rodded from end and the bundle removed
 Easier to clean and
 Be used for fouling liquid

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Since in our case the temperature difference is above 80 0
C and water is corrosive and fouling
liquid, I have selected these types of exchanger (i.e. split-ring rotating head exchanger). In short
it is for efficient and eases to clean.
3.4.1 Log Mean Temperature and True Mean Temperature
∆Tlm= [(T1-t2) -(T2-t1)]/ln[(T1-t2)/(T2-t1)] = [(106-83)- (87-53)]/ln [(106-83)/ (87-53)]
=64.80
C
R=(T1-T2)/(t2-t1) = (106-87) / (83-53) =0.633
S= (t2-t1)/ (T1-t1) = (83-53)/ (106-53) =0.566
Ft=0.89
∆Tm =0.89*64.80
C=57.670
C, Where
t1&t2, inlet and outlet temperature of the cold fluid
T1 &T2, inlet and outlet temperature of hot fluid
R and S, correction factor constant
∆Tlm, log mean temperature of the system and
∆Tm, true mean temperature of the system.
3.5 Heat Transfer Area (A0)
A0=Q/∆Tm*U= 84787.54w/(57.67k*1000w/m2
k = 1.47022m2

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3.6 Exchanger Layout and Tube Size
Iteration #1
Using a split-ring floating head exchanger for efficiency and ease of cleaning. Both fluids are
corrosive, the hot fluid has higher operating pressure than the cold one and higher temperature
than cold fluid, less viscous and has lower mass flow rate. Based on these facts I have made
economical decision by allocating the hottest fluid on the tube side and the cold fluid on the shell
side. Since the degree of corrosiveness of the fluid is not that much damaging, carbon steel can
be used for both shell and tubes construction purpose.
For the given area (calculated above), let start our calculation by using the minimum standard
tube diameter 19.05 mm (3/4 inch) outside diameter, 14.83 mm inside diameter, 1.83 m Long
tubes (the last minimum standard size) on a triangular (this is because it gives higher heat
transfer rates) 23.81 mm pitch (pitch=1.25do, it is recommended). The above minimum standard
tube dimension is selected based on the heat transfer area.
3.7 Numbers of Tube
Area of one tube (neglecting thickness of tube sheet)
A1=π*19.05*10-3
m*1.83m=0.10952m2
Numbers of tubes (Nt)
Nt =1.47022m2
/0.1095m2
=12.4, let say 12
So, for 4 passes, tubes per pass =3
Check the tube-side velocity at this stage to see if it looks reasonable.
Tube cross section area (Ac) =π/4(14.83*10-3
)2
=1.727*10-4
m2
Area per pass (ACP) =3*1.727* 10 -4
m2
=5.2* 10 -4
m2
Volume flow rate for hot fluid = 1353.6/954.708 = 3.94*10-4
m3
/s

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Tube side velocity, Ut=volume flow rate/area per pass= (3.94*10 -4
m3/s)/5.2* 10 -4
m2
=0.8m/s,
this is un-satisfactory (since it is not in the range of 1-2 m/s), it may be modified depending on
the result of tube side pressure drop.
3.8Tube Bundle Diameter and Shell Diameter (Db.)
From Table 12.4 of [2], for 4 tube passes,
K1=0.175
n1=2.285
Pt=1.25d0= 23.8125mm
Db. = d0* (Nt/k1)1/n1
=19.05(12/0.175)1/(2.285)
=121.2mm =0.1212m
For a split-ring floating head exchanger the typical shell clearance from Figure 12.10 of [2]is
48mm by extrapolating two point, so the shell inside diameter (Ds.),
Ds=Db+clerance=121.2+48 mm= 169.2mm =0.1692m
3.9 Tube-Side Heat Transfer Coefficient (hi)
Re = (ῤ*ut*di)/ µ= (960kg/m3
) *(0.8m/s) *(14.83*10-3
m)/ (0.9612kg/m.hr.) =4.1*104
, it is
turbulent flow
Pr = (Cp* µ)/kf = (4.214kJ/kg.k)*(0.9612kg/m.hr)/ (0.678w/m. k) =1.66
L/di=1830/14.83 = 123.4
From figure 12.23, jh =3.5*10-3
(hi*di)/kf=Nu=jh*Re*Pr0.33
= (3.5*10-3
) *(41000) *(1.66)0.33
= 170
hi=(Nu*kf)/di= (170*0.678w/m. k)/ (14.83*10-3
m) = 7772w/m2
k.

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3.10 Shell Side Heat Transfer Coefficient
 The triangular and rotated square patterns give higher heat-transfer rates, but at the
expense of a higher pressure drop than the square pattern. A square, or rotated square
arrangement, is used for heavily fouling fluids, where it is necessary to mechanically
clean the outside of the tubes. The recommended tube pitch (distance between tube
centers) is 1.25 times the tube outside diameter; and this will normally be used unless
process requirements dictate otherwise. Since the streams fluids are not heavily fouled I
have preferred the triangular pattern of the tube.
 The standard baffle spacing (LB) is in the range of (0.2 to 1) time the shell internal
diameter. A close baffle spacing will give higher heat transfer coefficients but at the
expanse of higher pressure drop. The optimum baffle spacing will usually be between 0.3
to 0.5 times of shell diameter. Thus for the cause of this problem I have taken 0.4 times
of shell diameter.
LB =0.4*Ds =68mm
Pt. (pitch) = 1.25=d0 = 23.8125mm
Calculate cross-flow area
Acs = (Pt – d0) *(Ds*lB)/Pt= (0.2) *(1.15*10 -2
m2
) = 2.3* 10 -3
m2
Equivalent diameter (hydraulic diameter)
De = 1.1/d0 (Pt
2
-0.9175d0
2
)= 13.52mm
Shell side velocity
Volume flow rate for feed = (0.41666kg/s)/957.9kg/m3
= 4.8 * 10 -4
m3
/s
Us = Vc/Acs= (4.8*10-4
m3
/s)/(2.3*10 -3
m2
) =0.25m/s (this velocity is almost satisfactory because
the optimum shell side velocity of liquid is in the range of 0.3 to 1m/s)
Re = (us*de* ρs)/µs= (947.3kg/m3
) *(0.25m/s) *(13.52*10-3
m)/ (1.607kg/m.hr)= 7.173*103
, it is
turbulent flow.
Pr = (Cp*µs)/kf = (4.004kJ/kg.k)*(1.607kg/m.hr.)/(0.61w/m. k)= 2.93

16
NB: Generally, a baffle cut of 20 to 25 per cent will be the optimum, giving good heat-transfer
rates, without excessive pressure drop. Let take 25% baffle cut as first trial. This must give good
heat transfer coefficient without too high pressure drop. Thus from figure 12.29, jh =7*10-3
Nu = (hs*de)/ kf =jh*Re*Pr0.33
= 126
hs =Nu*kf/de =5678w/m2
. K
3.11 Overall Heat Transfer Coefficient
D0/di =1.285
1/U0 = 1/h0+1/h0d+d0ln (d0/di)/2kw+d0/di (1/hid+1/hi)=8.7808 * 10-4
m20
C/w
U0 = 1138.4w/m20
C, this is above from my guess (but the deviation is 12.15% only it is
acceptable value). The numbers of tube possibly reduced but first let us check the pressure drop.
3.12 Pressure Drop
3.12.1 Tube Side
Jf = 3.5 * 10 -3
∆Pt = Nt *[8*jf (L/di) +2.5] *(ρ*ut
2
/2) =12*[8*3.5*10 -3
(1830/14.83) +2.5] *[960kg/m3
*
(0.8m/s) 2
/2] = 21.953k. Pa=0.22bar, it is in our range of assumption
3.12.2Shell Side Pressure Drop (∆Ps)
∆Ps = 8jf *(Ds/de) *(L/LB) *(ρ*us
2
)/2, from the figure jf = 5*10-3
= 4*4.5*10-3
(169.2/13.52) *
(1830/68) * (947.3kg/m3
) * (0.25m/s) 2
= 0.36kpa = 0.00366 bar, this in our range of assumption
Thus the work is accomplished, with slightly some errors. The result will be discussed on the
following sheet. The spread sheet is representing all the steps I have followed above and has
little correction of some values.

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Shell and TubeHeatExchangerDesign Spreadsheet
Project: ProjectNo.:
ItemNo.: - Service: By: Date/Time: 17-May-17 17:37
Tube Side Shell Ustart = 1000.00 W/m²°C
Hotwater frombottom Fluid Name Acetone solution Then the required transfer A = 1.470 m2
0.376 Flow (M), Kg/s 0.4 12
106 Temp. in, °C 53 4
87 Temp. out, °C 83 3 3.00
Av. Density 960 r, Kg/m3
947.3 0.00052 m²
Av. Viscosity 0.268 m, mNs/m2
0.446 0.000392 m³/s
Av. HeatCapacity 4.214 cp, kJ/kg°C 4.004 0.756212 m/s
HeatExchanged 85 Q, kW 85
Av. ThermalConductivity 0.6780 k, W/m°C 0.6100
Fouling Resistance 0.0002 R,m²°C/W 0.0002 Tube Pitch = 0.0238125 m
Pattern = Tri.
LMTD 64.8 °C 0.068 m
Corrected LMTD 57.7 °C 0.001818 m²
0.013526 m
0.000440 m³/s
Step 2. Inputtubing OD, BWGand Tube OD 0.0191 m 0.24 m/s
length (can be trialand error). BWG 14 6,950
Tube ID, d = 0.01483 m 3
Tube Length, L = 1.83 m 0.0070
Area ofone tube = 0.109 m² 70
0.00017 m² ho = 3,138 W/m²°C
R1 = 0.00032
R2 = 0.00020
page number 648 Bundle diameter = 0.121188 m R3 = 0.000053
Shelldiameter = 0.133688 m R4 = 0.0004
40232 Overallheattransfer coefficient= 1002.7 W/m²°C
1.66
123
page number 665 jh 0.0035
166.55
hi = 7614.48 W/m²°C
Step 6. Shellside heattransfer
coefficient
Area ofShell=
EquivalentDiameter, de =
PrandtlNo. =
Volumetricflowrate =
Baffle Spacing =
Number oftubesrequired =
No. ofpasses=
Shellside velocity =
Nusseltnumber =
Shellside ReynoldsNo., NRe =
Volumetricflow =
Av. Velocity =
Segmentalbaffle cut25% page number 673 jh =
Step. 4 Bundle and Shell
diameter
Step. 5 Tube side heat
transfer coefficient
Tubeside ReynoldsNo., NRe =
Crosssectionalarea oftube =
Preheater
Length / ID =
Step 1. Inputflows, conditionsand propertiesdata for shellside and tubeside
PrandtlNo. =
Nusseltnumber =
1
ABNET MENGESHA DUBE
Area oftubesper pass=
preheating of10% (w/w) ofacetone solution
Tubesper pass=
Step 3. Startconfiguring the exchanger. Begin with the assumed overallheattransfer
coefficientto thispoint:

18
ShellandTubeHeatExchangerDesignSpreadsheet
Project: ProjectNo.:
ItemNo.: - Service: By: Date/Time: 17-May-17 17:37 BWG/TubeWallThicknesses
8 0.165
9 0.148 0.017
10 0.134 0.014
Tube Side Shell Ustart= 1000.00 W/m²°C 11 0.120 0.014
Hotwaterfrombottom FluidName Acetonesolution ThentherequiredtransferA= 1.470 m2
12 0.109 0.011
0.376 Flow(M),Kg/s 0.4 12 13 0.095 0.014
106 Temp.in,°C 53 4 14 0.083 0.012
87 Temp.out,°C 83 3 3.00 15 0.072 0.011
Av.Density 960 r,Kg/m3
947.3 0.00052 m² 16 0.065 0.007
Av.Viscosity 0.268 m,mNs/m2
0.446 0.000392 m³/s
Av.HeatCapacity 4.214 cp,kJ/kg°C 4.004 0.756212 m/s
HeatExchanged 85 Q,kW 85
Av.ThermalConductivity 0.6780 k,W/m°C 0.6100
FoulingResistance 0.0002 R,m²°C/W 0.0002 TubePitch= 0.0238125 m
Pattern= Tri.
LMTD 64.8 °C 0.068 m
CorrectedLMTD 57.7 °C 0.001818 m²
0.013526 m
0.000440 m³/s
Step2. InputtubingOD,BWGand TubeOD 0.0191 m 0.24 m/s
length(canbetrialanderror). BWG 14 6,950
TubeID,d= 0.01483 m 3
TubeLength,L= 1.83 m 0.0070
Areaofonetube= 0.109 m² 70
0.00017 m² ho= 3,138 W/m²°C
R1= 0.00032 For2t.p For4t.p For6t.p
R2= 0.00020 For2tubepasses k1 0.249 0.175 0.0743
pagenumber648 Bundlediameter= 0.121188 m R3= 0.000053 n1 2.207 2.285 2.499
Shelldiameter= 0.133688 m R4= 0.0004 Clearance 0.0125
40232 Overallheattransfercoefficient= 1002.7 W/m²°C
1.66
123
pagenumber665 jh 0.0035
166.55
hi= 7614.48 W/m²°C
pagenumber646
pagenumber649
pagenumber649
Step6. Shellsideheattransfer
coefficient
AreaofShell=
EquivalentDiameter,de=
PrandtlNo.=
Volumetricflowrate=
BaffleSpacing=
Numberoftubesrequired=
No.ofpasses=
Shellsidevelocity=
Nusseltnumber=
ShellsideReynoldsNo.,NRe=
Volumetricflow=
Av.Velocity=
Segmentalbafflecut25%pagenumber673jh=
Step.4BundleandShell
diameter
Step.5Tubesideheat
transfercoefficient
TubesideReynoldsNo.,NRe=
Crosssectionalareaoftube=
Preheater
Length/ID=
Step1. Inputflows,conditionsandpropertiesdataforshellsideandtubeside
PrandtlNo.=
Nusseltnumber=
1
ABNETMENGESHADUBE
Areaoftubesperpass=
preheatingof10%(w/w)ofacetonesolution
Tubesperpass=
Step3. Startconfiguringtheexchanger. Beginwiththeassumedoverallheattransfer
coefficienttothispoint:

19
4 Summary of Notes Taken from the Literature
1 Tube dimensions
 Tube diameters in the range 5/8 in. (16 mm) to 2 in. (50 mm) are used. The smaller
diameters 5/8 to 1 in. (16 to 25 mm) are preferred for most duties, as they will give more
compact, and therefore cheaper, exchangers. Larger tubes are easier to clean by
mechanical methods and would be selected for heavily fouling fluids. Since our fluid is
not heavily fouled the smaller diameter is preferred.
 The tube thickness (gauge) is selected to withstand the internal pressure and give an
adequate corrosion allowance.
 The preferred lengths of tubes for heat exchangers are: 6 ft. (1.83 m), 8 ft (2.44 m), 12 ft
(3.66 m), 16 ft (4.88 m) 20 ft (6.10 m), 24 ft (7.32 m). For a given surface area, the use of
longer tubes will reduce the shell diameter; which will generally result in a lower cost
exchanger, particularly for high shell pressures. The optimum tube length to shell
diameter will usually fall within the range of 5 to 10.
2 Tube arrangements
The tubes in an exchange are usually arranged in an equilateral triangular, square, or rotated
square pattern; The triangular and rotated square patterns give higher heat-transfer rates, but at
the expense of a higher pressure drop than the square pattern. A square, or rotated square
arrangement, is used for heavily fouling fluids, where it is necessary to mechanically clean the
outside of the tubes. The recommended tube pitch (distance between tube centers) is 1.25 times
the tube outside diameter; and this will normally be used unless process requirements dictate
otherwise. In our case the fluids are not heavily fouling and triangular pitch arrangement is
preferred.

20
3 Fluid allocation
Where no phase change occurs, the following factors will determine the allocation of the fluid
streams to the shell or tubes.
Tube side fluid Shell side fluid Reasons
More corrosive Reduce the cost of expensive alloy or clad components.
More fouled Give better control over the design fluid velocity, and
the higher allowable velocity in the tubes will reduce
fouling. And the tube is easier to clean than shell
Higher
temperature fluid
Reduce the overall cost (when temperature is very high)
Reduce the shell surface temperatures, and hence the
need for lagging to reduce heat loss, or for safety
reasons. (when temperature is moderate )
Higher pressure
fluid
High-pressure tubes will be cheaper than a high-
pressure shell.
lowest allowable
pressure
drop fluid
higher heat-transfer coefficients will be
obtained on the tube-side than the shell-side
More viscous
fluid
A higher heat-transfer coefficient will be obtained by
allocating the more viscous material to the shell-side,
providing the flow is turbulent.
If turbulent flow cannot be achieved in the shell it is
better to place the fluid in the tubes.
Lowest flow
rate
Will normally give the most economical design.

21
4 Shell and tube fluid velocities
High velocities will give high heat-transfer coefficients but also a high-pressure drop. The
velocity must be high enough to prevent any suspended solids settling, but not so high as to
cause erosion. High velocities will reduce fouling. Typical design velocities for liquid (since all
our fluids are liquid) are given below:
 Tube-side, process fluids: 1 to 2 m/s, maximum 4 m/s if required to reduce fouling;
Water: 1.5 to 2.5 m/s.
 Shell-side: 0.3 to 1 m/s.
5 Stream temperatures
The closer the temperature approach used (the difference between the outlet temperatures of one
stream and the inlet temperature of the other stream) the larger will be the heat-transfer area
required for a given duty. The optimum value will depend on the application, and can only be
determined by making an economic analysis of alternative designs. As a general guide the
greater temperature difference should be at least 20 0
C. When the heat exchange is between
process fluids for heat recovery the optimum approach temperatures will normally not be lower
than 20 0C.
6 Pressure drop
In many applications the pressure drops available to drive the fluids through the exchanger will
be set by the process conditions, and the available pressure drop will vary from a few mill bars in
vacuum service to several bars in pressure systems. The values suggested below can be used as a
general guide, and will normally give designs that are near the optimum.
Liquids:
Viscosity <1 mN s/m2
35 kN/m2
(this range of viscosity and pressure drop)
1 to 10 mN s/m2
50 70 kN/m2

22
5 References
[1]. Indian Standard (IS: 4503-1967): Specification for Shell and Tube Type Heat Exchangers,
BIS 2007, New Delhi.
[2]. R. K. Sinnott, Coulson & Richardson’s Chemical Engineering: Chemical Engineering
Design (volume 6), Butterworth-Heinemann, 3 rd. ed. 1999.
[3]. D. Q. Kern, Process Heat Transfer, McGraw-Hill Book Company, Int. ed. 1965.
[4] Dutta B.K. „Heat Transfer-Principles and Applications‟, PHI Pvt. Ltd., New Delhi, 1st
ed.
2006.

23
PART II
DISTILLATION COLUMN DESIGN

24
5 Introduction
Distillation is a physical process for the separation of liquid mixtures that is based on differences
in the boiling points of the constituent components. Distillation is the most widely separation
process used in many industries.
Distillation applications
Distillation makes about 95% of all current industrial separation processes. It has been used in
chemical industries, pharmaceutical and food industries, and environmental technologies and in
petroleum-refineries.
The most common use is after a chemical reactor where we obtain some products. Distillation is
used in order to separate the desired product from the rest obtaining a high purity product.
As we can see, distillation is applied for many different processes because of its reliability,
simplicity and low-capital costs although these systems have relatively high energy consumption.
Distillation theory
The process of distillation begins with a feed stream that requires treatment. The feed is
separated into two fractions in a conventional column, the light product and the heavy product.
Throughout the report, the feed molar flow rate F will be reported by (kgmole/h) feed mole
fraction is z and the stage that feed enters is denoted by NF (normally the tray on which the
characteristics of the fluid is closest to that of the feed)
The liquid leaving the top of the column is the light component, while the liquid leaving the
bottom of the column is the heavy component. Liquid leaving the bottom of the column is split
into a bottoms product and a fraction that is made available for boiling. The re-boiler (heat
exchanger) is employed to boil the portion of the bottom liquid that is not drawn off as product.
The vapor produced flows up through the column and comes into intimate contact with the
down-flowing liquid. After the vapor reaches and leaves the top of the column, another heat
exchanger (the condenser) is encountered where heat is removed from the vapor to condensate it.
The condensed liquid is split into two streams. One is the overhead product; the other liquid
stream is called reflux and is returned to the top of the column to improve the separation.

25
Fundamental concepts
The vapor-liquid equilibrium on each stage is the central part of the distillation theory. The most
difficult part in the distillation column design is the description of this equilibrium between the
vapor and the liquid. To derive the following equations, we base on ideal mixtures so the vapor-
liquid equilibrium can be derived from Raoult´s law: The partial vapor pressure of a component
in a mixture is equal to the vapor pressure of the pure component at the temperature multiplied
by its mole fraction in the mixture.
For an ideal gas, according to Dalton´s law:
p𝑖= 𝑦𝑖 𝑃
Therefore, the equilibrium between the vapor and the liquid for ideal mixtures is:
Yi=xi (pi/P)
Note: However, we have assumed ideal equilibrium vapor-liquid; actually there are few systems
that work as ideal mixtures. For this reason, the simulations will be done with non-ideal
mixtures.
5.1 Problem Statement
1500kg/hr. of 10/90 (w/w) % acetone water mixture will be heated from 20 0
C to t2, to prepare
the mixture for the distillation process that produces top product of 90% (w/w) acetone and the
rest water, and the bottom products of (1/99) % (mole/mole) acetone/water. Design distillation
column for this service.
Note: here the designer must, determine column height and diameter, weir dimension, plate total
area, plate active area, calming zone area, plate hole dimension and pitch and material of
construction.

26
Assumptions/facts
Column efficiency of 50% and pressure drop per plate of 1.5 KPa may be assumed. We can take
the minimum liquid flow as 80% of the maximum rate both above and below the feed plate.
Feed temperature
Feed temperature is a major factor influencing the overall heat balance of a distillation column
system. Increments in the feed enthalpy can help reduce the required energy input from the re-
boiler at the same degree of separation. Installing a feed preheater is a very common process
option to minimize re-boiler heat duty.
If the feed preheater can be integrated with other valuable process streams (as a heating
medium), overall energy efficiency of the distillation system can be improved further. However,
increasing the feed temperature does not always improve the overall energy efficiency of a
distillation unit. Excessive feed temperature increments can cause a significant amount of flash
of heavy key and non-key components at the distillation column feed zone. In this case, a higher
amount of reflux stream is necessary to maintain required overhead distillate purities. This
augmented reflux ratio thus requires a higher boil-up ratio. Overall energy efficiency is
eventually aggravated. Therefore, careful review of the feed temperature and phase is critical to
minimize the overall energy consumption of the distillation unit.

27
Vapor-Liquid Equilibrium Data
Components
No. Formula Molar Mass Name
1 C3H6O 58.08 Acetone
2 H2O 18.015 Water
Constant Value
Pressure 101.33 KPa
Data Table
T [K] x1 [mole/mole] y1 [mole/mole]
368.25 0.00800 0.13800
363.25 0.01600 0.27700
355.25 0.03300 0.47900
349.35 0.05200 0.60400
345.35 0.07200 0.67500
342.65 0.09400 0.71900
340.75 0.11700 0.73800
338.15 0.17100 0.77600
336.45 0.23700 0.80000
335.15 0.31800 0.82200
334.15 0.42000 0.83900
333.05 0.55400 0.86300
331.35 0.73600 0.90900
(T - temperature, x - liquid mole fraction, y - vapor mole fraction

28
5.2 Mass Balance and Determination of No of Theoretical Stage
Table 6: Feed and product composition
Component Feed mole fraction Top product mole
fraction
Bottom product mole
fraction
Acetone 0.033333 0.7361 0.01
Water 0.96666 0.2639 0.99
Feed condition (q-line)
Bubble point saturated liquid feed
 L’
=L+F
 V=V’
The tope operating line (ROL)
Y= (L/V) *X+Xd /(R+1) ……………………………1
R=L/D
V=L+D
The striping section operating line
Y= ((VB+1)/VB) X-XB/VB………………………………….2
q= (L’-L)/ F=heat to vaporize 1 mole of feed/molar latent heat of feed
VB=L’/V’
L’=V’+B
Mass balance for total flow
F=D+B……………………………………….3
Mass balance for component flow
Xf*F=Xd *D+Xb*B…………………………………...4

29
Average molecular weight of the feed=0.033*58.08+0.9667*18 =19.3167kg/kmole
Feed flow rate= (1500kg/hr.)/19.3167kg/kmole =77.653kmole/hr.
From equation 3&4 we get
B=75.2kmole/hr. and
D=2.46kmole/hr.
For this problem the condition of minimum reflux occurs where the top operating line just
touches the equilibrium curve at the point where the q-line cuts the curve.
From the graph I have drown and cut the equilibrium curve where the q-line touch the curve, we
get xd/ (Rmin+1) =0.54
Rmin=0.36
As the flow above the feed point will be small, a high reflux ratio is justified; the condenser duty
will be small. Thus take R=3*Rmin =1.09
At this reflux ratio, phi =0.3522V=L+D and R=L/D………………………………………. (a)
From R and D values
L=2.6814kmole/hr. and V=5.1414kmole/hr.
Rectifying section slope = 0.455
Since our feed is at its bubble point temperature q=1 and
L’=F+L=80.3344kmole/hr. L’=B+V’ and V’=5.1344
Stripping section slope=L’/V’=S=7.09
Bubble point of feed (from the data shown in table) = 83°C, since our feed is at its bubble point
temperature, the heat required to vaporize 1 mole of the given feed and latent heat of the feed
almost equal.

30
I have preferred to calculate the number of stages than using the MC Cabe Thiele graph, since
the bottom (stripping section) is at a very low concentration for less volatile component. But the
top section need 2 stages this is from the graph.
Ns
x
=log [(k/s-1) (xr
’
/xb-1)/(1/s*(k-1)) +1] +1
Log (k/s)
From vapor liquid equilibrium data
K=y1/x1=0.17275/0.01=17.275
And xr=0.1
Ns
r
=5
Thus the total number of try is 2(above feed) +5(blow feed) =7
Number of real stages= (7-1)/0.5=12
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
VLED for acetone water mixture at 1atm
x1 mole fraction of acetone in liquid state
y1molefractionofacetoneinvaporstate

31
Table 7: Summary of material balance around the distillation column
Kmole/hr. for
top
Kg/hr. Kg/s Kmole/hr. for
bottom
Kg/hr. Kg/s
V=5.1414 279.9 0.0775 V’=5.1344 208.5 0.0579167
L=2.6814 127.37 0.0353806 L’=80.3344 1478.153 0.4106
D=2.46 116.85 0.03246 B=75.2 1383.7 0.38436
Data from McCabe Thiele diagram
Number of stages = 12
Slope of top operating line = 0.455
Slope of bottom operating line = 7.09
Top composition = 73.61 mole% acetone
Bottom composition = 1 mole% acetone (a)
Minimum reflux ratio = 0.36
5.3 Estimation of Physical Properties
Here it is important to know information about both the top and bottom of the column. Useful
information includes temperature, pressure, column pressure drop, densities, molecular weights,
surface tensions, and number of stages
Column top pressure= 101325 Pa (1 atm)
Column pressure drop=1.5*1000 ×12=18000Pa
Pressure drop of 1.5kPa per tray is specified

32
TOP SECTION BOTTOM SECTION
XD= 0.7361
Column top pressure= 101325 Pa (1.0147
bar) and temperature= 58.2 °C
𝜌 𝑣 =𝑃𝑀/𝑅𝑇=101325 ×54.433/331.35×8.314×103
= 2.002 kg/m3
𝜌𝑙 =787.06 kg/m3
(density of the mixture)
(water density= 983.8 and acetone
density=734.3779kg/m3
at 58.2 °C)
Average molecular weight of vapor:
M=54.433
Average molecular weight of liquid: M=47.5
Surface tension, 𝔡 =30.8254×10-3
N/m
Column bottom pressure=101325
+18000
= 121.33kPa (=1.1933bar)
Boiling point of water at (1.1933
bar) = 105 °C (bottom contains 99 moles
%
water)
From the steam table at 1.2133 and
105°C: 𝜌 𝑣= 0.705; 𝜌𝑙 = 955.11 kg/m3
Average molecular weight of vapor:
M=40.58
Average molecular weight of liquid:
M=18.4
𝔡 =58×10-3
N/m
5.4 Plat Spacing
The plate spacing will depend on the column diameter and operating conditions. Plate spacing
from 0.15 m to 1.0 m are typically used. The smaller the diameter, the smaller the spacing. Small
columns will use close spacing. Columns with diameters above 1.0 m, plate spacing of 0.3 m to
0.6 m are normally used. A good initial estimate is 0.5 m. thus I have used 500mm plate spacing
as first trail.
5.5 Column Diameter Estimation
Vapor and liquid flow rates will vary along the column, so plate design needs to be considered
both above and below the feed. Using plate spacing and FLV, you can obtain the value of K from
figure.

33
There is a range of vapor and liquid flow rates in which the column needs to be operated. Too
low or too high of rates can result in various inefficiencies in the column operation. For example,
if the vapor rate is too high, flooding will occur. However, it is not safe to operate on the
flooding line. Instead, columns are typically designed for 80% of flooding at the maximum flow
rate. Obtain a new velocity with this 80%, and use the velocity to calculate a maximum
volumetric flow rate. Using this and the velocity, we can calculate a net area necessary for vapor
flow through the plate. Also need to assume a down comer area. Now a column cross sectional
area can be calculated.
To calculate the column diameter an estimate of the net area Anis required. As a first trial take
the down comer area as 12 per cent of the total, and assume that the hole active area is 10 per
cent.
1st
trial is started with the following considerations:
 Design is performed for 80% flooding at maximum gas flow rate.
 Total down comer top and bottom seal area is 12% of the net area.
 The hole area to active area ratio as 15%
Flow parameter (FLV) based on mass flow rate,
Rectifying section slope*(ρv/ρl)0.5
0.455* (1.9463/786.3)0.5
= 0.02264
Capacity parameter
Csb=0.125 m/s
Correction for surface tensions
K1’=0.125 * (31/20)0.2
= 0.136451
Velocity@ flooding
Unf = k1’* ((ρl-ρv)/ρv)0.5
=2.702m/s
FLV = striping section slope
*(ρv/ρl)0.5
= 7.09*(0.705/955.11)0.5
= 0.2
Capacity parameter
Csb=0.11 m/s
Correction for surface tensions
K1’=0.11*(58/20)0.2
= 0.136
Flooding velocity = Unf = k1’*
((ρl-ρv)/ρv)0.5
= 5.004m/s
The linear design gas velocity

34
The linear design gas velocity (Un) based on
net area (80% flooding): U n = 0.8*2.702 m/s
= 2.162m/s
The maximum volumetric vapor flow rate
(Qmax) = (V*M)/ρv =
(279.85kg/hr.)/2.162kg/m3
= 0.03596m3
/s
Net area required:
Qmax/Un = 0.03596/2.162 = 0.016633m2
Totals tower cross-section area:
0.016633/0.88 = 0.019m2
(Total down comer top and bottom seal area is
12% of the net area)
Colum (tower) diameter= (0.019/0.7854)0.5
= 0.1552m
maximum allowable superficial vapor velocity
(based on total column area)
= Ѷ= (-0.171Lt
2
+0.27Lt-0.047) *((ρl-ρv)/ρv)1/2
= 0.8961m/s
Dc = √ ((4*Qmax)/ᴫ ρv Ѷ) = 0.126
(Un) based on net area (80%
flooding):
U n = 0.8*5.004 m/s = 4.003m/s
The maximum volumetric vapor flow rate
(Qmax) = (V*M)/ρv =
(208.5kg/hr.)/0.705kg/m3
= 0.0822m3
/s
Net area required:
Qmax/Un = 0.0822/4.003 = 0.020535m2
Totals tower cross-section area:
0.020535/0.88 =
0.0233m2
(Total down comer top and bottom seal
area is 12% of the net area)
Dc = (0.020535/0.7854)0.5
= 0.1617m
Use the higher value of the tower diameter for the uniformity between sections, if the difference
is not greater than 20%. In this case, the bottom diameter is used both in top and bottom sections.
0.1617m
To read FLV from the figure, the following must be restricted

35
1. Hole size less than 6.5 mm. Entrainment may be greater with larger hole sizes.
2. Weir height less than 15 per cent of the plate spacing.
3. Non-foaming systems.
4. Hole: active area ratio greater than 0.10;
5.6 Selection of Liquid-Flow Arrangement
Liquid volumetric flow rate in the top section = 127.37/3600 × 786.3 =4.5 * 10 -5
m3/s
Liquid volumetric flow rate in the bottom section = 1478.15/3600 × 955 = 4.3 * 10 – 4
m3/s
Therefore, single pass cross-flow sieve plate is chosen for this service.
5.7 Make Provisional Tray Layout
This step includes to calculate/select down-comer area, active area, perforated area, hole area and
size, weir height, weir length. It has done step by step
Figure 1: Provisional area and dimension
From the graph below, the ratio of down-comer area (Ad) to column cross-sectional area (Ac) can
be determined from the ratio of weir length (lw) to column diameter (Dc) and vice versa.
Figure 2: Relation between down comer area and weir length

36
Parameters Formula Values
Dc= tower diameter 0.1617m =0.162m
AT= total/cross sectional area ᴫ*Dc
2
/4 0.020536m2
AD =down comer area 0.12 AT 0.0024643 m2
An =net area AT-AD 0.018072 m2
AA =active area AT-2AD 0.0156074m2
Lw/Dc 0.77 from the figure
AD/AT 0.12
Lw =weir length 0.121275m
Ah =hole area 0.1 AA 0.00156074
Standard sizes for trays and good assumptions for the first iteration are: weir height, hw = 50mm;
hole diameter, Dh = 5mm; plate thickness, Pt = 5mm
5.8 Check the Weeping Rate
Here we compare the actual vapor velocity to the minimum vapor velocity, if velocity is too low
fluid will "weep" through the tray holes. If the weeping rate is unsatisfactory, return to step 6 and
choose different values for the plate layout dimensions. From the chart in step 4, it can be seen
that there is a minimum vapor flow rate below which the liquid "weeps" from the tray above.
For the remaining steps in this design process, it is recommended to check your assumptions
after each step and revise them as necessary in order to maintain operation in the "sweet spot" of
the vapor rate vs. liquid rate plot. Additional iterations may be required as you move through the
procedure.
Calculate the maximum liquid flow rate. Calculate the minimum liquid flow rate at 70%
turndown (recommended).

37
Maximum liquid flow rate (Lwc=mmax) =
0.0353806kg/s
Minimum liquid flow rate (mmin) =
0.7*0.0353806kg/s = 0.02477kg/s
Maximum weir crest, hwc = 750 * (Lwc/Lw
*ρl)2/3
= 750 * (0.0353806/
(0.121275*787.06))2/3
=3.87007mm liquid height
Minimum weir crest, hwc = 750 * (Lwc/Lw
*ρl)2/3
= 750 * (0.02477/
(0.121275*787.06))2/3
The constant (k2) of weep point correlation
= 30.2 at
hwc +hw = 50mm+3.0514mm = 53.0514mm
using minimum liquid flow rate
Minimum vapor velocity Umin at weep point
=(k2-0.9(25.4-Dh))/ ρl = (30.2-0.9(25.4-5)/
(2.008)1/2
) = 8.3555m/s
Actual minimum vapor velocity at minimum
vapor flow rate = (70% Qmax) AH
=(0.7*0.03596m3
/s)/(0.00156074m2
) = 16m/s
So minimum operating rate will be well above
Weep point
Maximum liquid flow rate (mmax) =
0.4106kg/s
Minimum liquid flow rate (mmin) =
0.7*0.4106kg/s = 0.28742kg/s
Maximum weir crest, hwc = 750 * (Lwc/Lw
*ρl)2/3
= 750 * (0.4106/ (0.121275*955.11))2/3
Minimum weir crest, hwc = 750 * (Lwc/Lw
*ρl)2/3
= 750 * (0.28742/
(0.121275*955.11))2/3
For this section k2 is 30.4 at
hwc +hw = 50mm+13.74642mm=63.75mm
Minimum vapor velocity Umin at weep point
=(k2-0.9(25.4-Dh))/ ρl = (30.4-0.9(25.4-5)/
(0.705)1/2
) = 14.34m/s
Actual minimum vapor velocity at minimum
vapor flow rate = (70% Qmax) AH
=(0.7*0.0822m3
/s)/(0.00156074m2
) = 37m/s
So minimum operating rate will be well above
Weep point

38
Figure 3: Weep-point correlation
5.9 Check Plate Pressure Drop
Maximum vapor velocity through a hole
Umax= Qmax/AH =0.03596/0.00156074
= 23m/s
Maximum dry plate pressure drop hd =
51*(Umax/Co)2
* (ρv/ρl) =51*(23/0.84)2
*
(2.002/787.06) = 98mm
Co = 0.84 for AH/AP =15% and pt/hd =1
Residual head hr = (12.5 *103
)/ρl = 16mm
The total plate drop
ht =hd +(hw + hwc) +hr = 98+(50+3.05) +16 =
177mm
Maximum vapor velocity through a hole
Umax= Qmax/AH =0.0822/0.00156074
= 53m/s
Maximum dry plate pressure drop hd =
51*(Umax/Co)2
* (ρv/ρl) =51*(53/0.84)2
*
(0.705/955.11) = 108mm
Co = 0.84 for AH/AP =15% and pt/hd =1
Residual head hr = (12.5 *103
)/ρl = 13mm
ht =hd +(hw + hwc) +hr = 108+(50+13.75) +13
= 185mm
At first trial 1.5kpa pressure drop was assumed and according to the above pressure drop
calculation it is acceptable.

39
5. 10 Down-Comer Backup Liquid and Down-Comer Residence
Time
5.10.1Down-comer design [back-up]
The down comer area and plate spacing must be such that the level of the liquid and frothin the
down comer is well below the top of the outlet weir on the plate above. If the level rises above
the outlet weir the column will flood.
The back-up of liquid in the down comer is caused by the pressure drop over the plate(the down
comer in effect forms one leg of a U-tube) and the resistance to flow in the down comer itself;
In terms of clear liquid, the down comer back-up is given by:
hb =(hw+hcw) +ht+hdc
Where hb = down comer back-up, measured from plate surface, mm,
hdc = head loss in the down comer, mm. it is estimated as =166(Lwd/ρl*Am)2
Where Lwd =liquid flow rate in down comer, kg/s,
Am = either the down comer area Ad or the clearance area under the down comer
Aap; whichever is the smaller, m2
.
The clearance area under the down comer is given by:
Aap = hap *lw
Where hap is height of the bottom edge of the apron above the plate. This height is
Normally set at 5 to 10 mm below the outlet weir height: hap = hw – (5 to 10 mm), let take 10mm
as first trial. Thus hap = 40mm and Lw = 121.275mm
Aap = hap *lw 121.275*40 = 0.00485m2
and AD = 0.0024643m2
. Thus Ad is smaller of the two.

40
hdc =166(Lwd/ρl*Am)2
=
166((0.0353806kg/s)/787kg/m3
*0.0024643m2
)
=0.005525mm
hb = (hwc +hw) +ht+hdc =
3.87+50+112.7+0.00553 = 167mm
Down comer residence time (tdrt)
=(AD*hb*ρl)/Lwd
=(787*0.167*0.0024643)/0.0353806 =9.2s
hdc =166(Lwd/ρl*Am)2
=
166((0.4106kg/s)/0.0024643m2
*
955.11kg/m3
)2
= 5.052mm
hb = (hwc +hw) +ht+hdc =
13.74+50+143+5.053 = 212mm
Down comer residence time (tdrt)
=(AD*hb*ρl)/Lwd
=(955.11*0.212*0.0024643)/0.4106 =12.2s
0.5*(plate spacing + weir height) = ½ (500+50) = 275mm> hb no flooding.
Therefore, the plate spacing and down comer residence time in both section meet the design
requirement.
5.11 Plate Layout
Perforated area (Ap) = AA – Acz - AES
Where, Acz = calming zone area, AES = area occupied by edge strip
Lw/DT=0.77; now, Өc= 1000
, Angle subtended by the chord (edge plate), = 1800
-1000
= 800
The unperforated edge strip (edge plate) mean length from the geometry:
LES = (DT- 50mm) (ᴫ *80)/180 = (0.1617m -0.05m) *3.143*80/180 =0.156m
AES = 0.005*LES = 0.05*0.156 =0.0058m2
Use 50mm wide calming zones. The approximate mean length of zones:
Lcz =Weir length (Lw) + Width of un perforated edge strip =0.121275+0.05=0.171275 m
Acz = 2(0.005*Lcz) = 0.0017128 m2
Therefore, Perforated area (Ap) = AA – Acz - AES =0.0156074 -0.0090628=0.01152m2
Ah/Ap = 0.00156074/0.01152 =0.1355

41
From figure Lp/Dh = 2.7 it is satisfactory. Since the optimum rage is between 2.5 to 4.0
For equilateral triangle pitch
Ah/Ap =0.1355= 0.9(Dh/Lp)2
, (0.9/0.1355)1/2
= Lp/Dh =2.67And Lp = 5mm*2.7 = 13.5mm
Number of hole
 Area of one hole = 1.94 * 10-5
m2
 Number of hole = 0.0156074/1.94*10-5
= 795
5.12 Flooding and Entrainment Checking
Actual vapor velocity based on the net area
(An),
Uv=Qmax/An=(0.03596m3
/s)/0.018072m2
=1.99m
/s
%flooding = Uv/Unf *100 = (1.99/2.7) *100
=74%
The fractional entrainment at FLV = 0.02264 and
actual flooding velocity of 74% is ψ =0.085
Effects of ψ on murphree plate efficiency can be
estimated from: Ea = (Emv)/ (1+(ψ Emv/1-ψ))
=0.478
Emv = 0.5 (murphree vapor efficiency 50%) and
Ea = 0.478 (murphree plate efficiency corrected
from entrainment )
Actual vapor velocity based on the net area
(An),
Uv=Qmax/An=(0.0822m3
/s)/0.018072m2
=4.2m
/s
%flooding = Uv/Unf *100 = (4.4/5.2) *100
=80.8%
The fractional entrainment at FLV = 0.2 and
actual flooding velocity of 82% is ψ =0.009
Effects of ψ on murphree plate efficiency can
be estimated from: Ea = (Emv)/ (1+(ψ Emv/1-
ψ)) =0.498
Emv = 0.5 (murphree vapor efficiency 50%)
and
Ea = 0.498 (murphree plate efficiency
corrected from entrainment )
Thus the actual flooding is almost below the design flooding value 80%. And usually ψ <0.1 is
desirable.

42
6 Conclusion
In order to have stable operation in a distillation column, the vapor and liquid flows must be
managed. Requirements are:
• Vapor should flow only through the open regions of the tray between the down comers
• Liquid should flow only through the down comers
• Liquid should not weep through tray perforations
• Liquid should not be carried up the column entrained in the vapor
• Vapor should not be carried down the column in the liquid
• Vapor should not bubble up through the down comers
These requirements can be met if the column is properly sized and the tray layouts correctly
determined.
Tray layout and column internal design is quite specialized, so final designs are usually done by
specialists; however, it is common for preliminary designs to be done by ordinarily superhuman
process engineers and students. This project work is intended to give you an overview of how
this can be done by the student of chemical engineering, who take apparatus design course, so
that it won't be a complete mystery when you have to do it for your detailed design project.
Basically in order to get a preliminary sizing for our column, we need to obtain values for
• The tray efficiency
• The column diameter
• The pressure drop
• The column height

43
7 Reference
1. Robert E. Treybal, Mass Transfer Operations, McGraw-Hill, Inc., 3rd
ed. 1981.
2. Perry’s Chemical Engineers’ Handbook, McGraw-Hill, Inc., 8th
ed. 1997.
3. R. K. Sinnott, Coulson & Richardson’s Chemical Engineering: Chemical
Engineering Design (vol. 6), Butterworth-Heinemann, 3rd
ed. 1999.
4. Perry’s Chemical Engineers’ Handbook, McGraw-Hill Companies, 7th
ed. 1997.
5. Henry Z. Kister, Distillation Design, McGraw-Hill, Inc., 1st ed. 1992.

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