INDIVIDUAL DESIGN

ASSIGNMENT 1 – CHEMICAL DESIGN

(Design of a de – ethanizer column)

Name: Perera A. T. K.

Index Number: 100375D

Date of Submission: 05/01/2015

INTRODUCTION

Ethylene is a widely using industrial chemical as a raw material in producing various

products like polyethylene, polyvinyl chloride, ethylene dichloride and etc. Ethylene is

mostly produced by steam or catalytic cracking of Naphtha or Natural Gas. As per the

economists views the global demand for ethylene may increase at a compounding rate of 10

million tons per year. In Sri Lankan scenario, amount of ethylene based products imported in

year 2011 was 52,000 MT. This figure can be expected to be grown by now with the

increasing global demand. So to match this increasing demand in house ethylene production

in the country can be considered.

As the comprehensive design project of our group, it was decided to address this matter and

develop an ethylene production facility with steam cracking of Natural Gas as there’s a huge

potential of having Natural Gas in Mannar basin oil exploration sites. However to begin the

process imported LNG is used. The process flow for the production follows several

separation columns before and after the steam cracker along with other common equipments

like heat exchangers and compressors. For separation of the key components along the

process, distillation can be used in most of the steps.

The mixture of hydrocarbons obtained after the cracking should be separated at several stages

and separation of C2 components from C3 and other heavier hydrocarbon components is

crucial prior to the final fractionation of ethylene from C2 components. This separation is

done using the de-ethanizer, a distillation column designed to match the separation

requirements.

Based on the boiling points at standard temperature and pressure C2 components can be

separated from C3 and other heavier components using distillation.

Table 1.1 : Boiling points of Feed components at 1 atm

Component Boiling Point at 1 atm (°C)

Ethylene -104

Ethane -88

Ethyne -84

Propylene -48

Propane -42

Propyne -23.2

Butene -6

Butane 0

Butyne 8.1

Pyrolysis gasoline 32 - 204

Cyclobutadiene 41.2

Benzene 80.1

Toluene 110.6

Styrene 145

Pyrolysis fuel oil 169 – 579.4

Maximum recovery of C2 components at the top product is required as the final product

outcome is ethylene and ethyne (Acetylene) can be converted into ethylene by catalytic

hydrogenation which is financially beneficial. At the final fractionators ethylene is separated

from ethane and ethane is sent back to steam cracker for production of ethylene.

Commercially the separation of C2 components is done almost completely from C3

components. Because of these reasons maximum recovery of 99.99% of ethyne at the

distillate and 0.01% recovery of propene at the top product are taken as design considerations

while designing the tower secondary de-ethanizer.

Further for the conversion purpose of ethyne, the top product is collected and transferred to

the next unit in gaseous phase. Therefore a partial condenser is used at the top of the

distillation tower.

DISTILLATION COLUMN

Distillation columns can be designed with different column internals. According to the

column internal the column may be a packed column or a tray column. In both columns there

are few advantages and disadvantages. Therefore the selection of the tower type should be

done carefully.

Packed bed columns

These are columns in which the internal is packed with random or structured shape packing

material and provides a larger surface area for vapor liquid mixing. These are most often used

for absorption but can be used in distillation too. The associated advantages and

disadvantages of packed bed columns can be listed as follows.

Advantages:

For columns less than 0.6 m in diameter, most efficient tower type.

Lower pressure drop compared to tray columns.

Can be used to handle corrosive liquid mixtures as packing material can be made by inert

material.

Suitable for thermally sensitive liquid separations.

Disadvantages:

Packing material may get damaged during installation or under extreme temperatures due to

thermal expansion or contraction.

Contact efficiency is low when liquid flow rates are smaller.

Higher cost at high liquid flow rates.

Tray distillation columns

These are columns in which the internal upward and downward flow of liquid and vapor is

disturbed by placing trays with a definite gap in between them in order to obtain a good

liquid vapor mixing. These are widely used in distillation applications. The number of trays

varies according to the application. This type also has some pros and cons.

Advantages:

Most efficient type of distillation column when column diameter exceeds 0.6 m.

Suitable for cryogenic distillation applications as it is easy to equip with cooling coils.

Liquid vapor contact in cross flow tray column is more effective than countercurrent flow in

packed columns.

Can use for high liquid flow rates cost effectively.

Can handle cryogenic conditions of distillation.

Disadvantages

Higher pressure drop compared to packed beds.

Possibility of foaming as a result of agitation of liquid through vapor.

Considering all the above factors, it is clear that packed bed columns are not suitable for

operation at cryogenic conditions and hence can be eliminated from selection. Then the

viable option becomes using tray type column for the separation of desired components. This

is because the separation should be done under cryogenic conditions.

Selection of tray type for the column

The trays that are used as column internals can be designed in several ways. The types

include sieve trays, valve trays and bubble cap trays. Each of these types brings its own

advantages and disadvantages to the distillation column. Those advantages and disadvantages

can be considered as follows.

Sieve tray Valve tray Bubble cap tray

Cost Ratio (Compared

to Sieve tray)

1 1.5 3

Efficiency Very similar for all three types

Capacity Very similar for all three types

Pressure drop Lowest value Moderate Highest value

Application Depend on the vapor

flow rate

Weeping can occur

Has a greater

flexibility

Suitable for low

vapor flow rate

applications as

design provides a

liquid seal.

Considering all the factors above discussed it is decided to select valve type trays to complete

the column internal as it has a greater flexibility and the capital cost for the construction also

is reasonable when compared to bubble cap trays. Further the capacities and efficiencies

being nearly similar for all three types it would be better to select a cost effective and flexible

design for the column.

Flow arrangement

The arrangement of flow on the trays can be decided

MATERIAL BALANCE

Since the feed consists of number of components, multicomponent distillation should be

considered in order to separate the components as desired. Therefore determination of key

components for the material balance becomes a crucial step of calculation. As the major

requirement of de-ethanizer is to separate C2 components from C3 and other heavier

components, ethyne was selected to be the light key with a recovery of 99.99 mol% at the

distillate and propene was selected to be the heavy key with a recovery of 0.01 mol% at the

distillate.

Light key: The most volatile component in the bottoms, but in a significant

concentration is known as the light key. Therefore more volatile components than

light key does not go to the bottoms

Heavy key: The least volatile component in the distillate, but in a significant

concentration is known as the heavy key. Therefore less volatile components than the

heavy key does not go the distillate

Only ethyne (Light key) and propene (heavy key) is distributed between top and bottom

products as there are no other components in between the light and heavy key according to

boiling points.

The following formulae were used in calculation of mass fractions, mole fractions and molar

or mass flow rates during material balance.

During calculations a constant molar overflow of components through the column was

assumed.

Based on the above formulae the calculated feed composition for the secondary de-ethanizer

is as follows.

Table 2.1: Feed composition for the secondary de-ethanizer

General equation for the material balance;

For a multicomponent distillation column, the material balance equations can be expressed in

the following form.

Overall material balance for the tower;

Overall material balance for the ith

component;

Where;

Component MT/day W/W % kmol/day mol/mol % MW (kg/kmol)

Ethylene 294.3411 57.3304 10512.1821 59.9835 28

Ethane 193.4509 37.6795 6448.3633 36.7950 30

Ethyne 4.1544 0.8092 159.7846 0.9117 26

Propylene 5.5164 1.0745 131.3429 0.7495 42

Propane 0.8196 0.1596 18.6273 0.1063 44

Propyne 0.1195 0.0233 2.9875 0.0170 40

Butene 0.8955 0.1744 15.9911 0.0912 56

Butane 1.2167 0.2370 20.9776 0.1197 58

Butyne 9.1517 1.7825 169.4759 0.9670 54

Pyrolysis gasoline 1.3393 0.2609 13.3930 0.0764 100

Cyclobutadiene 0.3328 0.0648 6.4000 0.0365 52

Benzene 1.4555 0.2835 18.6603 0.1065 78

Toluene 0.224 0.0436 2.4348 0.0139 92

Styrene 0.0568 0.0111 1.1360 0.0065 50

Pyrolysis fuel oil 0.3375 0.0657 3.3750 0.0193 100

Total 513.4117 100 17525.1313 100

F = Feed flow rate

D = Distillate flow rate

W = Bottom flow rate

XiF = Molar fraction of ith

component in the Feed

XiD = Molar fraction of ith component in the Distillate

XiW = Molar fraction of ith

component in the Bottom

Further, XiF, XiD and XiW can be calculated using the following formulae.

Being heavy key and light key only ethyne and propene are distributed between top and

bottom products according to the specified recovery molar fractions. The components lighter

than light key are only found in tracer amounts in the bottoms and components heavier than

heavy key are only found in tracer amounts in the top product. Therefore the material balance

for the distillation tower will be as follows.

D kmol/day

F kmol/day

W kmol/day

Table 2.2: Distribution of feed composition over distillate and bottom

Therefore;

Component XiF

F. XiF

kmol/day

D. XiD

(kmol/day)

W. XiW

(kmol/day)

XiD XiW

Ethylene 0.5998 10512.1821 10512.1821 0 0.6140 0.0000

Ethane 0.3679 6448.3633 6448.3633 0 0.3766 0.0000

Ethyne 0.0091 159.7846 159.7686 0.0160 0.0093 0.0000

Propylene 0.0075 131.3429 0.0131 131.3297 0.0000 0.3244

Propane 0.0011 18.6273 0 18.6273 0.0000 0.0460

Propyne 0.0002 2.9875 0 2.9875 0.0000 0.0074

Butene 0.0009 15.9911 0 15.9911 0.0000 0.0395

Butane 0.0012 20.9776 0 20.9776 0.0000 0.0518

Butyne 0.0097 169.4759 0 169.4759 0.0000 0.4187

Pyrolysis gasoline 0.0008 13.3930 0 13.3930 0.0000 0.0331

Cyclobutadiene 0.0004 6.4000 0 6.4000 0.0000 0.0158

Benzene 0.0011 18.6603 0 18.6603 0.0000 0.0461

Toluene 0.0001 2.4348 0 2.4348 0.0000 0.0060

Styrene 0.0001 1.1360 0 1.1360 0.0000 0.0028

Pyrolysis fuel oil 0.0002 3.3750 0 3.3750 0.0000 0.0083

Total 1 17525.1313 17120.3272 404.8041 1 1

Further for flow rate calculations inside the tower material balances for rectifying and

stripping section are required.

Rectification Section Stripping Section

Assuming Constant Molar overflow,

Material balance for the system shown, Material Balance for the system shown,

V= L+D - - - (eqn 2.8) L’ = V’+W - - - (eqn 2.10)

Further, R = L/D [defined] - - - (eqn 2.9)

Material Balance over feed plate,

F+ L+ V’ = V+L’

V’ – V = L’ – L – F - - - (eqn 2.11)

Further, as feed is saturated liquid,

L’ = L+ F - - - (eqn 2.12)

D

L

L V

W

L’ V’

F

L’ V’

L V F

ENERGY BALANCE

Energy balance for a distillation column simply implies the condensing and heating energy

requirements in condenser and reboiler. In order to find these it is essential to know distillate

and bottom products temperatures. This is because the energy required for condensation of

distillate or vaporization of bottoms is a function of its enthalpy at that temperature.

Temperature Calculations for Towers

Bubble point calculation

A temperature is assumed for the bubble point of tower

Feed or liquid phase composition and operating pressure are known

K values are obtained from the literature for the corresponding temperature and

pressure for each component.

Vapor phase mole fractions were calculated for all the components using,

Where;

Yi = mole fraction of ith

component in vapor phase

Ki = K value of ith

component where,

Here, = Saturated vapor pressure of i

th component and,

P = Operating pressure of the distillation tower

Xi = mole fraction of ith

component in liquid phase

If ƩYi = 1 the assumed temperature is correct and if not a new K value is calculated

for one component and a temperature that satisfies the new K value is taken.

Calculation was repeated from step 1 until ƩYi =1

Dew point calculation

The known composition is taken as vapor phase composition when operating

pressure is known.

A temperature is assumed for the dew point and K values for all the components

were found at corresponding temperature.

Liquid phase composition (Xi) was calculated for all the composition using the

following equation.

If ƩXi = 1 the assumed temperature is correct and if not a new K value for one

component is calculated.

New temperature (Tnew) to match the new K value is chosen and K values for all

the components at new temperature are found.

Calculation was repeated until ƩXi =1.

The calculated liquid and vapor phase compositions and temperatures of distillate and bottom

can be tabulated as below.

Table 2.3: Distillate and Bottom temperatures with compositions

Distillate temperature (Vapor

dew point)

Bottom temperature ( Liquid

bubble point)

Component Yi K (-23.15) Xi2 Xi K(95.85) Yi

Ethylene 0.6140 1.3117 0.468101 0.0000 0

Ethane 0.3766 0.725 0.519517 0.0000 0

Ethyne 0.0093 0.75 0.012443 0.0000 0

Propylene 0.0000 0 0.3244 2.431493 0.788844

Propane 0.0000 0 0.0460 2.2 0.101234

Propyne 0.0000 0 0.0074 0.6285 0.004638

Butene 0.0000 0 0.0395 0.148 0.005846

Butane 0.0000 0 0.0518 0.835 0.043271

Butyne 0.0000 0 0.4187 0.1342 0.056184

Pyrolysis gasoline 0.0000 0 0.0331 0.0283 0.000936

Cyclobutadiene 0.0000 0 0.0158 0.2034 0.003216

Benzene 0.0000 0 0.0461 0.0639 0.002946

Toluene 0.0000 0 0.0060 0.0127 7.64E-05

Styrene 0.0000 0 0.0028 0.00637 1.79E-05

Pyrolysis fuel oil 0.0000 0 0.0083 0.0283 0.000236

Total 1.0000 1.00006 1.0000 1.007447

Therefore;

Distillate temperature = -23.15 °C

Bottom temperature = 95.85 °C

Average column temperature =

Therefore, average column temperature = 36.35 °C

Since the distillate is taken as saturated vapor and bottom product is taken as saturated liquid,

only latent heats of vaporization and condensation counts for the energy balance or Reboiler

and condenser heat loads.

Condenser Heat Load

Where, λL = Latent heat of liquid mixture which is at equilibrium with distillate vapor

L = Liquid flow rate of rectifying section (only this part is condensed while product is taken

as vapor)

Latent heat of a mixture can be calculated as,

From the calculations, λL = 487052.212 J/kg

LW = Liquid mass flow rate in rectifying section = 224622.1158 kg/day

Therefore,

Condenser Heat Load = 109402.698 MJ

Reboiler Heat Load

In similar manner, Reboiler heat load is calculated. Here Latent heat of vaporizing vapor and

its mass flow rate is taken for calculation.

From the calculations, λL = 503139.3 J/kg

LW = Vapor mass flow rate in stripping section = 929639.3296 kg/day

Therefore,

Condenser Heat Load = 467738.0815 MJ

COLUMN DESIGN

Minimum Reflux Ratio (Rm)

Minimum reflux ratio for the column can be found using Underwood equations for multi

component systems.

First Underwood equation;

Second Underwood equation;

Where;

αi = Relative volatility of ith

component at average column temperature

ϕ = A factor defined for the calculation

q = L/F where, L- Liquid fraction of feed and F- Feed flow rate

q = q factor of the feed

Table 2.4: α values of feed components at column average temperature

Component α value at 36.35 °C

Ethylene 5.2941

Ethane 3.4118

Ethyne 4.0588

Propylene 1

Propane 0.8235

Propyne 0.3981

Butene 0.1071

Butane 0.2368

Butyne 0.0908

Pyrolysis gasoline 0.0164

Cyclobutadiene 0.1255

Benzene 0.0198

Toluene 0.0074

Styrene 0.0017

Pyrolysis fuel oil 0.0164

α values are defined with respect to heavy key of the system which in this case is propylene.

Where; KA = K value of considering component

KB = K value of heavy key

q Factor for the feed = L/F,

As the feed is saturated liquid or liquid at its boiling point L=F,

Therefore, q = F/F =1

For the convenience of calculation, the components below the heavy key and which are not included

in the mixture with great percentages were considered as groups.

C3 components are considered as propane (pink), C4 components as butyne (green) and other heavier

components as Heptene (Blue) as most of them are C7 components.

Table 2.5: Feed composition to the tower

Component XiF F. XiF kmol/day

Ethylene 0.5998 10512.1821

Ethane 0.3679 6448.3633

Ethyne 0.0091 159.7846

Propylene 0.0075 131.3429

Propane 0.0011 18.6273

Propyne 0.0002 2.9875

Butene 0.0009 15.9911

Butane 0.0012 20.9776

Butyne 0.0097 169.4759

Pyrolysis gasoline 0.0008 13.3930

Cyclobutadiene 0.0004 6.4000

Benzene 0.0011 18.6603

Toluene 0.0001 2.4348

Table 2.6: Categorized feed composition and their relative volatility at average temperature

Component XiF F. XiF kmol/day α Values

Ethylene 0.5998 10512.1821

5.2941

Ethane 0.3679 6448.3633

3.4118

Ethyne 0.0091 159.7846

4.0588

Propylene 0.0075 131.3429

1

Propane 0.0013 21.6148 0.8235

Butyne 0.0118 206.4446 0.0908

Heptene 0.0027 42.0241 0.0164

Using 2nd

Underwood equation the following polynomial equation is obtained.

Using MATLAB software, the roots of the above polynomial are founded to be,

4.0096 + 0.0446i, 4.0096 - 0.0446i, 1.001, 0.8467, 0.1007 and 0.0202

Choose ϕ = 1.001 [In between heavy key and light key]

Then, applying ϕ = 1.001 in first Underwood equation,

Rm = 0.3007

Optimum Reflux ratio (R)

The optimum reflux ratio of distillation lies between 1.2 -1.5 times Rm at many instances.

Therefore, choose

Using equations from eqn 2.9 to eqn 2.12, When R = 0.45,

Styrene 0.0001 1.1360

Pyrolysis fuel oil 0.0002 3.3750

Total 1 17525.1313

L = 7704. 1472 kmol/day

V =24824.2945 kmol/day

L’ = 25229.2785 kmol/day

V’ = 24824.4744 kmol/day

All flow rates within the column are positive. Therefore the Reflux Ratio of 0.45 is acceptable.

Minimum Number of theoretical stages (Nm)

Fenske equation for determination of Nm,

Where;

αlk ,ave = relative volatility of light key ( the geometric average)

Then, Minimum number of theoretical stages = 10.68

Number of theoretical stages required for the separation (N)

Gilliland correlation for the estimation of number of theoretical plates is used,

Fitted equations

x = a finite value for the ratio between reflux ratios

Here, x= 0.103

For, 0.01 ≤ x ≤ 0.9

So the number of theoretical stages required,

N= 22.9119 ≈ 23

Feed tray location

Feed tray location at minimum number of theoretical plates (NF, min) is given by;

α values are given at column average temperature, and molar fractions at total reflux are similar to

calculated values as there are no distributed components in between the heavy key and light key.

Therefore; from calculation;

NF, min = 6.3846

Feed tray location at finite Reflux (NF)

Feed tray Location when R=0.45,

NF = 13.697

Efficiency of column (E)

O’connel correlation provide with a means of obtaining overall column efficiency. This correlation is

mostly suitable for hydrocarbon mixtures. Therefore can be used for this system also with a high

accuracy.

To read the graph, the product of Molar average liquid viscosity (µa mNs/m2) and average relative

volatility of light key (αa) is required.

Viscosity of a liquid mixture can be expressed with following relationship,

For molar fractions of components in liquid mixture, molar fractions of feed components are used.

Viscosity of liquid mixture = µ

αa = 4.8415 [from calculation of Nm]

From the graph of O’connel correlation,

Overall efficiency of the column = 55%

This lies between typical efficiency range of 30 -70%. Therefore it is an acceptable value.

Actual Number of plates (Na)

The overall column efficiency (E) is related to Number of ideal stages (N) as per the

following equation.

From solving the equation,

Actual number of plates in the column (Na) =41.658 ≈ 42

Diameter Calculation

Mass flow rates

Rectifying section

Where;

Lw, Vw = liquid or vapor mass flow rate through rectifying section

L, V = Molar flow rates of liquid and vapor

ML,V,avg = Average molar mass of liquid or vapor

ML,avg = 29.156 kg/kmol, MV,avg = 29.734 kg/kmol

Mi = Molar mass of ith

component

Xi = Molar fraction of ith

component in liquid or vapor respectively

Therefore;

LW = 224622.116 kg/day = 2.5998 kg/s

VW = 738125.5727 kg/day = 8.5431 kg/s

Stripping section

Lw’, Vw

’ = liquid or vapor mass flow rate through stripping section

L’, V

’ = Molar flow rates of liquid and vapor

ML,avg = 41.1609 kg/kmol, MV,avg = 37.4485 kg/kmol

Therefore;

LW ‘ = 1038459.809 kg/day = 12.0192 kg/s

VW’ = 929639.3296 kg/day = 10.7597 kg/s

Average densities

Rectifying section

ρ L,avg = Average density of liquid phase in rectifying section = 556.0817 kg/m3

ρi = Density of ith

component in liquid phase

For vapor phase consider all the components behave as ideal gases and use ideal gas law.

Here, P = Operating pressure of the tower = 2000 kPa

R= Universal Gas Constant = 8.314 J/mol K

T = Average temperature in Kelvin = 252.575 K

ρV,avg = Average vapor density

ρV,avg

Stripping section

Here also same equations are used. But the notification is given as ( )’ for the convenience of

identification.

Therefore;

ρ L,avg’ = Average density of liquid phase in stripping section = 605.1199 kg/m

3

T = Average temperature in Kelvin = 312.075 K

ρV,avg’

Flow parameters

Flow parameters for rectifying and stripping section can be found using following equation.

When calculations for stripping section are done ( )’ is used as notification.

Rectifying section

Therefore, FLV = 0.0687

Stripping section

Then from calculation, FLV’ = 0.244

Area calculation

Rectifying section

From the graph for the relationship of Flow parameter vs. Flooding vapor velocity, (Vol 6,

p567)

K1 factor = 0.101 , for a tray spacing of 0.6 m.

Then,

Uf = Flooding vapor velocity based on the net column cross sectional area in m/s

Uf = 0.4360 m/s

Assume 70% of flooding condition inside the tower

Then, Actual vapor velocity (Ua ) can be given by,

Ua =0.3052 m/s

The volumetric flow rate of vapor can be given as,

Vg = 0.3017 m3/s

The net area of the column cross section can be then found using,

The down comer area of the column is usually,

The net area and column area is related by,

Therefore;

And the Active area of the plate is given by,

Where;

An = Net area of column cross section

AC = Column cross sectional area

Ad = Down comer area

Aa = Active area

Also the hole area of plate (Ah) can be found as.

From the calculation, following results are obtained;

An = 0.9885 m2

AC = 1.1233 m2

Ad = 0.1348 m2

Aa = 0.8537 m2

Ah = 0.08537 m2

Stripping section

Same procedure of calculation for rectifying section was used. But the notification is used as

( )’ for the convenience of identification.

Therefore the calculated parameter values are as follows,

From the graph for the relationship of Flow parameter vs. Flooding vapor velocity, (Vol 6,

p567)

K1’ factor = 0.076, for a tray spacing of 0.6 m.

Uf’ = 0.3396 m/s

Assume 70% of flooding is maintained in the tower

Then, Ua’ = 0.2377 m/s

Vg’ = 0.3727 m3/s

An’ = 1.568 m2

AC’ = 1.7818 m2

Ad’ = 0.2138 m2

Aa’ = 1.3542 m2

Ah’ = 0.1354 m2

Column diameters

Rectifying section

Once the column cross sectional area is known, the diameter of the column can be calculated

as follows.

Where,

DC = Column diameter

The column diameter for the rectifying section,

DC = 1.1959 m

Diameter of the column is greater than 1 m. Therefore the tray spacing of 0.6 m is acceptable.

Stripping section

Using the same equation, the diameter of the stripping section (DC’) is found to be,

DC’ = 1.5062 m

The diameter of the stripping section is greater than 1 m. Therefore the tray spacing of 0.6 m

is acceptable.

Column Height

Column height can be obtained by,

H = Column height

Na = Number of actual trays = 42

Tray spacing = 0.6 m

Therefore, H = 25.2 m