Assignment 1 – Chemical Design

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 Cost Ratio (Compared 1

Valve tray

Bubble cap tray

1.5

3

to Sieve tray) Efficiency

Very similar for all three types

Capacity

Very similar for all three types

Pressure drop

Lowest value

Application

Depend on the vapor Has flow rate

Moderate a

flexibility

Weeping can occur

Highest value greater Suitable vapor

for

low

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 Component

MT/day

W/W %

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

513.4117

100

17525.1313

100

Total

kmol/day mol/mol % MW (kg/kmol)

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;

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.

D kmol/day

F kmol/day

W kmol/day

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.

Table 2.2: Distribution of feed composition over distillate and bottom

Component Ethylene Ethane Ethyne Propylene Propane Propyne Butene Butane Butyne Pyrolysis gasoline Cyclobutadiene Benzene Toluene Styrene Pyrolysis fuel oil Total

Therefore;

W. XiW (kmol/day)

XiD

XiW

F. XiF kmol/day

D. XiD (kmol/day)

0.5998

10512.1821

10512.1821

0

0.6140

0.0000

0.3679

6448.3633

6448.3633

0

0.3766

0.0000

0.0091

159.7846

159.7686

0.0160

0.0093

0.0000

0.0075

131.3429

0.0131

131.3297

0.0000

0.3244

0.0011

18.6273

0

18.6273

0.0000

0.0460

0.0002

2.9875

0

2.9875

0.0000

0.0074

0.0009

15.9911

0

15.9911

0.0000

0.0395

0.0012

20.9776

0

20.9776

0.0000

0.0518

0.0097

169.4759

0

169.4759

0.0000

0.4187

0.0008

13.3930

0

13.3930

0.0000

0.0331

0.0004

6.4000

0

6.4000

0.0000

0.0158

0.0011

18.6603

0

18.6603

0.0000

0.0461

0.0001

2.4348

0

2.4348

0.0000

0.0060

0.0001

1.1360

0

1.1360

0.0000

0.0028

0.0002

3.3750

0

3.3750

0.0000

0.0083

1

17525.1313

17120.3272

404.8041

1

1

XiF

Further for flow rate calculations inside the tower material balances for rectifying and stripping section are required. Rectification Section

Stripping Section

F

D

V’ L V

L’

L W

Assuming Constant Molar overflow, Material balance for the system shown,

Material Balance for the system shown, L’ = V’+W - - - (eqn 2.10)

V= L+D - - - (eqn 2.8) Further, R = L/D [defined] - - - (eqn 2.9) Material Balance over feed plate, F+ L+ V’ = V+L’ F

V

L

V’ – V = L’ – L – F - - - (eqn 2.11) V’

L’

Further, as feed is saturated liquid, L’ = L+ F - - - (eqn 2.12)

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 ith 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 ƩX i =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

Component

Distillate temperature (Vapor dew point) Yi K (-23.15) Xi2

Ethylene Ethane Ethyne Propylene Propane Propyne Butene Butane Butyne Pyrolysis gasoline Cyclobutadiene Benzene Toluene Styrene Pyrolysis fuel oil Total

0.6140 0.3766 0.0093 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 1.0000

Therefore;

1.3117 0.725 0.75

0.468101 0.519517 0.012443 0 0 0 0 0 0 0 0 0 0 0 0 1.00006

Bottom temperature ( Liquid bubble point) Xi K(95.85) Yi 0.0000 0.0000 0.0000 0.3244 0.0460 0.0074 0.0395 0.0518 0.4187 0.0331 0.0158 0.0461 0.0060 0.0028 0.0083 1.0000

2.431493 2.2 0.6285 0.148 0.835 0.1342 0.0283 0.2034 0.0639 0.0127 0.00637 0.0283

0 0 0 0.788844 0.101234 0.004638 0.005846 0.043271 0.056184 0.000936 0.003216 0.002946 7.64E-05 1.79E-05 0.000236 1.007447

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 Ethylene Ethane Ethyne Propylene Propane Propyne Butene Butane Butyne Pyrolysis gasoline Cyclobutadiene Benzene Toluene Styrene Pyrolysis fuel oil

α value at 36.35 °C 5.2941 3.4118 4.0588 1 0.8235 0.3981 0.1071 0.2368 0.0908 0.0164 0.1255 0.0198 0.0074 0.0017 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

Styrene

0.0001

1.1360

Pyrolysis fuel oil

0.0002

3.3750

1

17525.1313

Total

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

Component

XiF

Ethylene

F. XiF kmol/day

α Values

Ethyne

0.0091

10512.1821 5.2941 3.4118 6448.3633 159.7846 4.0588

Propylene Propane

0.0075 0.0013

131.3429 1 21.6148 0.8235

Butyne

0.0118

206.4446 0.0908

Heptene

0.0027

42.0241 0.0164

Ethane

0.5998 0.3679

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,

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 (µ a mNs/m2) and average relative volatility of light key (α a) 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/m3 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