Gas to Liquids Plant Design

GAS TO LIQUIDS PLANT DESIGN

Prepared For

Dr. J.D. PALMER

Professor

Chemical Engineering Department

Louisiana Tech University

By

Bajracharya Prabuddha

Enaohwo Enakeme

Obe Akeem

Cmen 432 Students

May 5nd, 2011

1

EXECUTIVE SUMMARYThis report discusses the details of how we designed a Fischer-Tropsch Reaction unit (FTR) to convert syngas

(CO and H2) to hydrocarbons. The purpose of the design is to produce liquid fuels and reduce our over-

dependence on importation of oil from foreign countries.

In achieving our goal of effectively designing a Fischer-Tropsch Reaction unit to produce liquid fuel, design

objectives were adhered to strictly. The design was made to ensure that it does not pose any environmental,

health or safety hazards by ensuring that appropriate control valves were placed in needed areas. Rigorous

research was also done on the material of construction of the equipment to ensure that no reaction between the

components and the material of construction occur. This may lead to an unwanted product that could greatly

affect the purity and specification of our final product.

Also, rigorous economic analysis was done to ensure that the design is feasible. In the economic analysis, the

designed capital investment, expense costs of the air separation unit, syngas unit, FTR unit, and hydro-

isomerization unit were reflected. After an economic analysis on the design process, a net present value of 4.96

billion dollars was achieved under the process conditions defined in the report.

2

INTRODUCTIONSynthetic fuels production is one of the promising alternatives to petroleum fuels. Synthetic fuels can be

produced by converting natural gas to liquid petroleum products. One of these gases to liquid processes is

Fischer-Tropsch synthesis. Fischer-Tropsch is one of the oldest and mature processes to produce alternative fuel

source. This process converts carbon monoxide and hydrogen gas in combination called synthetic gas into

liquid hydrocarbon fuels like diesel and jet fuel.

The objective of this project is to provide preliminary design package for a grass roots Fischer-Tropsch reactor

unit with reactor effluent separation facilities. We are to design a safe, environmentally clean, thermally

integrated FTR unit operating cost utilization. The project should also reflect that the optimal cost/benefit

balance is achieved.

Some of the design specifications are:

Effective heat integration

Environmental, health or safety hazards properly mitigated

No continuous flaring/venting of hydrocarbons

The economic constraints for the plant are:

15 year project life, depreciation is SL over 15 years.

Tax is 33%

3% projected yearly inflation

Total capital investment estimated by multiplying equipment cost by 4.8

Operating expenses above and beyond utilities is estimated using 3% of the total capital investment

FTR unit will have 1 month turnaround to coincide with catalyst replacement

3

SUMMARYThe effective and efficient production of liquid fuel from syngas involves three processes. These processes are

conducted in the syngas unit (SU), Fischer-Tropsch Reaction (FTR) unit, and the hydro-isomerization (HI) unit.

In the syngas unit (SU), methane gas, steam, carbon dioxide, and oxygen are reacted in the syngas unit to

produce carbon monoxide and hydrogen gas. The syngas unit was designed to convert 500 MSCF/D (Actual

conditions: 500 PSIG, 100F) of clean methane to syngas. In the syngas unit, three reactions occur. The first

reaction is a steam reforming reaction between methane gas and water to produce syngas; the second reaction is

a partial oxidation reaction between methane gas and oxygen to produce carbon monoxide, and water. It should

be noted that the heat required to drive the endothermic auto thermal reforming reaction is supplied in-situ by

partial oxidation of methane.

In the Fischer-Tropsch reaction unit, the carbon monoxide and hydrogen gas from the syngas unit are fed into

two tubular fixed reactors arranged in series. In these reactors, different hydrocarbons were formed at specific

temperatures and pressures. The products from the FTR unit are sent to distillation columns to separate products

based on differences on their boiling points.

In the hydro-isomerization unit, the distillate and heavier fractions (material with greater than 350 F boiling

point) are fed to a catalytic hydro-isomerization reactor, where paraffin’s are isomerized and wax is converted

to lighter products. A catalyst was used to reduce the activation energy of the reaction. It should be noted that

this catalyst is sensitive to water and carbon monoxide; therefore the liquid feed to the hydro-isomerization

reactor was free from water above the solubility limit and the makeup gas has a carbon monoxide content no

greater than 0.1 mol%.

In order to check the viability of the design, an economic analysis is done to obtain a net present value that

would justify the investment to this process.

4

DISCUSSIONSELECTION OF THERMODYNAMIC MODEL

To ensure a good process simulation, an accurate thermodynamic model must be chosen. Choosing an accurate

thermodynamic model will depend on the components of the incoming feed stream and the products.

In order to simulate portions of the GTRL plant, we have to have an appropriate thermodynamic model. Choosing a

thermodynamic model will depend on the data available and the most cumbersome compounds. We consulted a decision

tree to narrow down the choice of thermodynamic models and we came up with SRK, and PR or their variants. We were

able to find data for different binary system and we used that data to compare with the data given by chemcad model. To

identify the best fit model that simulated the data we used sum of least square method and found out that MSRK produced

the best fit. The following diagrams show the graphical fitting for one of our component system

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 160

80

100

120

140

160

180

200

n-Butane/n-Heptane TP-xy Experimental

Exp LiquidExp vapor

n-Butane Mole Fraction

Temp

eratu

re C

In order to find the best thermodynamic model, the TP-XY for n-butane and n-heptane were plotted for SRK, PR, PSRK,

and MSRK. They were plotted alongside the experimental data and a regression analysis was done on the data to

determine what model was the best fit. The plots are shown below:

5

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 160

80

100

120

140

160

180

200

SRK/TP-xy Experimental

Exp LiquidExp vaporSRK liquidSRK vapor


Tem

pera

ture

C

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 160

80

100

120

140

160

180

200

Peng Robinson/TP-xy Experimental

Exp LiquidExp vaporPR LiquidPRvapor


Temp

eratu

re C

6

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 160

80

100

120

140

160

180

200

PSRK/TP-xy Experimental

Exp LiquidExp vaporPSRK liquidPSRK vapor


Tem

pera

ture

C

The graphs all seem similar and we can decide to just go with one of them, but to achieve even more accuracy in our

simulation, a regression analysis is done on the data for the vapor phase:

SRK PR PSRK MSRK0.700000000000001

0.750000000000001

0.800000000000001

0.850000000000001

0.900000000000001

0.950000000000001

1

R2 value vs. Thermodynamic Models

Hexane Vs HeptaneMethane vs. DecaneMethane vs. ButaneButane vs Heptane

Thermodynamic Models

R2 V

alue

Based on the analysis, we found out that some were better than others for different binary systems but for a global

thermodynamic model we found out that MSRK was the right choice. So MSRK was chosen as our thermodynamic

model.

7

MATERIAL OF CONSTRUCTIONWhen choosing the material of construction for any process, the corrosive effects of the components in the process must

be factored in, because it make the difference between having to replace equipment in the future and not having to do that

at all. The material of construction for the different equipment’s will depend on the contents of the feed and exit streams.

In order to choose the material of construction for our process, the NACE corrosion database was consulted. We were able

to find corrosion data for each major component:

Column1 CH4 Column2 Column3 Column4

# exposure medium Conc % temp (F) Corrosion rate(mpy)1 CS 100 25-225 <22 SS 12Cr 100 25-475 <23 Al alloy 100 25-225 <24 SS 304. 100 25-475 <25 SS 316. 100 25-475 <26 (Ni-30Mo) alloy 100 25-225 <2

Column1 Steam Column2 Column3 Column4# exposure medium Conc % temp (F) Corrosion rate(mpy)1 CS 100 212-475 <22 SS 12Cr 100 212-475 <23 Al alloy 100 212-475 <204 SS 304. 100 212-475 <25 SS 316. 100 212-475 <26 (Ni-30Mo) alloy 100 212-475 <2

Column1 H2 Column2 Column3 Column4


Column1 CO Column2 Column3 Column4


8

Column1 C2H6 Column2 Column3 Column4# exposure medium Conc % temp (F) Corrosion rate(mpy)1 CS 100 25-500 <22 SS 12Cr 100 25-500 <23 Al alloy 100 25-255 <24 SS 304. 100 25-500 <25 SS 316. 100 25-500 <26 (Ni-30Mo) alloy 100 25-500 <2

The corrosion data for C3 is similar to that of C4, so it is not shown.

Column1 C4H10 Column2 Column3 Column4


Column1 Naphtha (C5-C10) Column2 Column3 Column4


Column1 C12H26 Column2 Column3 Column4

# exposure medium Conc % temp (F) Corrosion rate(mpy)1 CS 100 25-225 <202 SS 12Cr 100 25-225 <203 Al alloy 100 25-225 <204 SS 304. 100 25-225 <205 SS 316. 100 25-225 <206 (Ni-30Mo) alloy 100 25-225 <27 tantalum 100 25-225 <2

9

The table below shows a suitable material for the syngas unit reactor, but not the corrosion rate.

Column1 CO2/N2/O2/H20 Column2 Column3 Column4

# exposure medium Conc % temp (F) Corrosion rate(mpy)

1 (chlorendic unsaturated polyester, glass fiber reinforced) 100 120

Based on the data collected in the tables above, a combination of materials has to be used for the process. For example, we

can use carbon steel for the naphtha reboiler, we also have to use a Nickel alloy for the equipment that will be separating

the diesel from naphtha. These are just examples of how the equipment will be chosen; safety handbooks will also be

consulted to see if the contents of the entry stream have any adverse reactions with our chosen material.

EQUIPMENT TYPE AND SIZING

10

In this FT synthesis processes utilized different equipment like heat exchangers, distillation columns. The brief

description of general equipment sizing and detailed description of specialized equipment sizing is shown

below.

1. Heat Exchangers : There are multiple heat exchangers integrated in our design .All of the heat

exchangers were sized using the following equation

Q=UA ∆T lm

Where “Q” is the duty, “U” is the overall heat transfer coefficient, “A” is the heat transfer area and ∆Tlm

is the log mean temperature difference. The overall heat transfer coefficient for the condenser is 850

W/m2°C, and for the reboiler is 1140 W/m2°C.

2. Distillation columns: Three distillation columns are used to meet the purity specifications of the

products. The tower height can be calculated with the help of number of trays and tray spacing .The

number of trays can be changed as the tower is optimized.

Therefore, the tower height (H) is given by,

H=¿of trays∗tray spacing

Volume of the tower is given by

V=π r2h

3. Reflux tank: The reflux tank was sized according to hold uptime and volumetric flow rate. The holdup

time for the reflux tanks was 10 min.

Reflux Tanks can be sized using the following relation:

V=2∗M∗tρ

Where V is the volume of the drum, M is the mass flow rate; t is the holdup time and ρis the liquid

density.

Diameter of the drum for totally condensed liquid is given by4 the following relation which considers

diameter to be 4 times the height.

D=[Vπ ]1 /3

4. Pumps: Centrifugal pumps were used because they are cost efficient and can be easily controlled using

controlled valves. Each of the pump was sized according the formula below;11

Power (Hp )=(Flow(gpm)∗∆ P( psi))/1714 /efficiency

5. Tray efficiency

Based on 41 sets of performance data for different sets of trays Drickamer and Bradford correlated the

overall stage efficiency. The data was collected for mainly hydrocarbons mixtures and few water and

miscible organics4.

The correlation produces the following empherical equation

Eo=13.3−66.8 log(μ)

Where μ is the viscosity of liquid in centipoise, Eo is the overall efficiency in percent.

SPECIALIZED EQUIPMENT DESIGN

Syngas Reactor12

Syngas reactor was one of the major equipment for our design .It is the reactor where methane feed with

addition of carbon dioxide, oxygen and steam is converted into carbon monoxide and hydrogen. For the syngas

reactor different reactions were given in the problem statement and are as follows:

CH 4+H 2O↔CO+3 H2

CO+H 2O↔CO2+H2

CH 4+32O

2→CO+2H 2

Here the first reaction is the partial oxidation reaction. This was modeled offline and the two reactions that

follow are equilibrium reaction these reactions were modeled in chemcad in an equilibrium reactor. The partial

oxidation reaction is exothermic and two equilibrium reactions are exothermic. The reactors were aligned in

series in such a way that the exothermic partial oxidation would provide the heat needed for the endothermic

equilibrium reaction. The reactor had to be optimized in such a way that the steam to methane ratio would be

0.5 mol/mol so as to prevent coking and hydrogen to carbon monoxide molar ratio out of the syngas reactor was

to be optimal at 2:1 molar ratio. The following optimization was done to achieve this.

0 500 1000 1500 2000 2500 3000 3500 4000 45001.95

22.052.1

2.152.2

2.25H2/CO ratio 1000ºF

preheat500ºF

Molar Flow rate CO2 (kmol/h)

H2/C

O m

olar

ratio

13

From the optimization we found out that the preheat temperature did not make any difference. The best ratio

was achieved at 2.01:1 molar ratio of hydrogen to carbon monoxide. The molar flow rates of components are

given in the table.

Best ratio at 3300 and 3400

T. (ºF) CO (kmol/hr) H2(kmol/hr)

500 8664 17410

750 8664 17410

1000 8664 17410

The above diagram shows the PFD of the syngas section. Methane, steam and carbon dioxide are sent in as

feed. These components are then sent to a preheater which heats the feed to 600 °F .The preheated feed is then

mixed with oxygen, compressed and sent to the syngas reactor. The product that comes out of the syngas reactor

is mostly carbon monoxide and hydrogen. The stream also contains some water and unreacted carbon dioxide

and oxygen. The stream coming out of the syngas reactor is then cooled by a train of heat exchangers and sent

to a phase separator. In the phase separator water is separated out as liquid and sent to waste treatment whereas

the gas is sent to FTR section of the plant.

14

FISHER TROPSCH REACTOR DESIGN

The FTR reactor was the most important piece of equipment in the project. Therefore a proper design of this

equipment was necessary. It was also one of the most difficult parts of the design. FTR reactor is the reactor

where conversion of syngas to different hydrocarbon chains takes place. The product selectivity of different

hydrocarbon chains was based on different equation as given below:

Conversion of syngas

xH 2+CO→H 2O+(CH2)nH 2

Product Selectivity

SC H4=r CH 4

−rCO=0.03∗T 3

SCn=0.04∗SCH 4n=2,3,4

The selectivity of C5+ was based on Anderson Shulz Flory probability distribution

W n

n=[ (1−α )2

α ]∗α n

The calculations for the FTR reactor had to be done by simultaneous solving of material, hydraulic and heat

balance around the reactor. The following equations were used in solving the reactor.

The material balance was given by the general mole balance around a packed bed reactor.

dXdW

=−r A

FAO∗ρb

Where the rate equation was given by Langmuir Hinshelwood form

−r A=k∗T 1∗ph2∗pco

(1+k2∗T 2∗pco )2

The rate equation here was heavily based on the partial pressures of the components and the temperature.

The hydraulic balance around the reactor was achieved using the Ergun equation relating catalyst weight to the

pressure drop.

dPdW

=−α∗T∗Po∗FT

2∗T o∗PPo

∗F¿

The heat balance around the reactor was achieved using:

dTdW

=

U∗aρ b

∗(T a−T )+r A∗∆ H rxn

FAO(Σθi∗C pi+∆C pX )

The overall heat transfer coefficient was given in the problem statement as:

15

U=0.385∗G0.8

D0.2

Where G is the inlet mass velocity and D is the tube diameter.

The equations above were solved simultaneously using math lab and we found out that the optimum conversion

was 30% and the pressure drop was 38.0 psi. The temperature profile down the reactor seemed to increase at the

end of the reactor and was calculated to be 650 K.

As seen in the figure above the conversion profile down the reactor increased until it reached a plateau.

Different boiler feed water temperatures were used to achieve optimum conversion. The optimum conversion

achieved was 30% at boiler feed water temperature of 450 K and catalyst weight of 45000 lbm. Above BFW

temperature of 450 K the conversion did increase but was unstable and led to a runaway reaction.

The figure above shows the temperature profile down the length of the reactor related by catalyst weight. The

conversion profile showed us that the temperature increased as the catalyst weight/length down the reactor

16

increased and around 50000 lbm of catalyst there was a runaway reaction. Therefore use of 45000 lbm of

catalyst which was below the runaway conditions was used and optimized.

The following summarizes the different values that we got for the FTR reactor.

Conversion (X) = 30%

Pressure Drop (ΔP) = 38.9 psi

GHSV = 100/hr

11 reactors in parallel in 1st train

8 reactors in parallel in 2nd train

1.4 in diameter tubes

1 in tube spacing

3857 tubes per reactor.

The figure above is the PFD of the FTR section. Syngas containing carbon monoxide, hydrogen and carbon

dioxide enters the FTR section from Syngas section. The stream is first sent to an absorber where carbon

dioxide is absorbed off using MEA. From the absorber we have almost pure carbon monoxide and hydrogen

coming out. This pure syngas is first sent to a heat exchanger where it is heated and then to a compressor. From

the compressor the syngas is then sent to the FTR reactor where conversions into different hydrocarbons take

place. The hydrocarbons out of the reactor are then sent to a heat exchanger and separation section for further

processing.

17

PRODUCT SELECTIVITY

After calculating the reactor conversion, we were able to calculate the product selectivity for the different

alkanes. Methane is dependent on temperature over the range of 390-450oF, the equation is below:

SC H4=rC H 4

−rCO=0.03T 3(

molsmol

)

T 3=exp [−10000( 1T

− 1473 )]

T is the reactor operating temperature in degree K.

After calculating the selectivity of methane, we could calculate the selectivity of the C2-C4 alkanes. They are

linked to that of methane by the following equations:

SCn=0.04 SCH 4(molmol )n=2,3,4.

The selectivity for C5+ can be calculated using the Anderson –Shultz-Flory equation:

M n=(1−α )α n−1

Mn is the relative mol fraction of carbon n.

α is the chain growth parameter, which is slightly temperature dependent and can be calculated using the

following formula:

α=0.93T 4

T 4=exp [250( 1T

− 1473 )]

Using these equations, the relative mole fraction of the hydrocarbons in the feed was calculated, and the actual

molar flow rates are gotten by multiplying the relative mole fractions by the converted molar flow rate of CO

into the system.

18

ABSORBER SIZING

Number of Equilibrium Stages

Absorption is a process in which a gas mixture is contacted with a liquid (the absorbent or solvent) to

selectively dissolves one or more components by mass transfer from the liquid. In our project two absorption

columns were used in different stages of our design. In the first stage of our design, an absorption column was

used to remove CH4 from the incoming gas stream. In the second stage of our design, an absorption column was

used to remove CO, H2, and CH4.

1 st Absorption column

From Seader’s Separation process principles,

fractionof solute absorbed= AN+1−AAN+1−1

Where,

A is the absorption factor; N is the number of equilibrium trays

The absorption factor is given as,

A= LKV

Where,

L is the volumetric flow rate of solute-free absorbent; V is volumetric flow rate of solute-free gas; K is the

vapor-liquid equilibrium ratio.

The incoming gas stream has the following composition.

Components Molar flow rate (kmol/h)

CO2 5.98

CH4 0.26

H2O 0.20

H2 53.71

O2 5.27

N2 0.18

CO 26.07

19

Total(V) 91.7

L=137.51kmolh

The K values of the different components were obtained at a pressure of 1.88 atm and a temperature of 25°C.

Components K-values

CO2 0.83

CO 0.023

H2 0.019

N2 0.015

O2 0.032

CH4 4.45

Components A-values

CO2 1.80

CO 64.54

H2 78.66

N2 100.54

O2 47.17

CH4 0.34

The fraction of each component that is absorbed at different number of stages is presented in a table below.

Components 1 2 3 4

CO2 0.64 0.83 0.92 0.96

CO 0.98 1.00 1.00 1.00

H2 0.99 1.00 1.00 1.00

N2 0.99 1.00 1.00 1.00

O2 0.98 1.00 1.00 1.00

CH4 0.25 0.31 0.33 0.33

20

The number of equilibrium stages needed in the first absorption column is 4

2 nd Absorption Column

T = 95 °F

P = 342.43 psia

Components kmol/h K A

H2 6.124 27.77 0.05402

CO 11.5 15.38 0.09753

H2O 0.1011

CH4 0.4399 6.65 0.22556

C2H5 0.0176 1.64 0.91463

C3H7 0.0176 0.584 2.56849

N-Butane 0.0176 0.195 7.69231

N-Pentane 0.7639 0.0713 21.0379

N-Hexane 0.71032 0.02 75

N-Heptane 0.66049 0.008 187.5

N-Octane 0.61416 0.0001 15000

N-Nonane 0.57107 0.00001 150000

N-Decane 0.531 0.000001 1500000

N-Undecane 0.4938 0.0000001 1.5E+07

N-Dodecane 0.45911 0.0000001 1.5E+07

N-Tridecane 0.4269 0.0001 15000

N-Tetradecane 0.397 1E-10 1.5E+10

N-Pentadecane 0.3691 1E-11 1.5E+11

N-Hexadecane 0.3432 1E-12 1.5E+12

N-Heptadecane 0.31913 1E-13 1.5E+13

N-Octadene 0.2967 1E-14 1.5E+14

N-Nonadecane 0.2759 1E-15 1.5E+15

N-Eicosane 0.2566 1E-16 1.5E+16

Total 25.7062

21

The fraction of each component that is absorbed at different number of stages is presented in a table below

Components 1 2 3 4 5 6 7 8

H2 0.05 0.05 0.05 0.05 0.05 0.05 0.05 0.05

CO 0.09 0.10 0.10 0.10 0.10 0.10 0.10 0.10

CH4 0.18 0.22 0.22 0.23 0.23 0.23 0.23 0.23

C2H5 0.48 0.64 0.72 0.76 0.79 0.82 0.83 0.85

C3H7 0.72 0.90 0.96 0.99 0.99 1.00 1.00 1.00

N-Butane 0.88 0.99 1.00 1.00 1.00 1.00 1.00 1.00

N-Pentane 0.95 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Hexane 0.99 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Heptane 0.99 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Octane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Nonane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Decane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Undecane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Dodecane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Tridecane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Tetradecane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Pentadecane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Hexadecane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Heptadecane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Octadene 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Nonadecane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

N-Eicosane 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

The number of equilibrium stages needed in the first absorption column is 8.

Overall Efficiency

Using the correlation by O’Connell (Seader), the overall efficiency is given as;

log Eo=1.597−0.199 log( KM LµL

ρ )−0.0896( log( K M LµL

ρ ))2

22

Where,

Ki = K-value of species being absorbed

ML =molecular weight of the liquid, lb./lbmol

µL = viscosity of the liquid, cP

ρL = density of the liquid, lb/ft3

1st Absorption column

Component Overall Stage Efficiency, %

CO 50.06

H2 50.85

2nd Absorption column

Component Overall Stage Efficiency, %

CO 6.83

H2 7.49

Tower Diameter

1 st Absorption Column

The flood velocity, Uf is given as;

U f=C ( ρL−ρVρV

)0.5

Where,

C = capacity parameter of Souders and Brown.

ρL = density of the absorbing liquid = 1000 kg/m3

ρV = density of the gas stream = 10.5 kg/m3

C=FST FF FHACF

Where,

FST = surface tension factor = (σ/20)0.5 =1.87

FF = foaming factor

FHA = 1.0 for Ah/Aa >/0.10 and 5(Ah/Aa) + 0.5 for 0.06 < Ah/Aa <0.1

σ = liquid surface tension = 70 dyne/cm

For no foaming systems, FF = 1.0; for many absorbers, FF may be 0.75 or even less (Seader).

23

FLV=( LM L

V MV)( ρV

ρL )0.5

=( 137.51∗1891.67∗13.97 )∗( 10.5

1000 )0.5

=0.1981

From Fig 6.24(Seader), with a tray spacing of 24inches, CF = 0.33ft/s

Because FLV < 0.1, Ad/A = 0.1 and FHA = 1.0. Thus,

C=(1.87 ) (0.75 ) (1 ) (0.33 )=0.46 fts

U f=0.46 (1000−10.510.5 )

0.5

=4.47 fts

DT=( 4V MV

f U f π (1− Ad

A ) ρV )0.5

=(4∗91.67

3600 ∗13.97

0.80∗( 4.473.28 )∗3.14∗0.9∗10.5 )

0.5

=0.66m

2 nd Absorption Column

FLV=( LM L

V MV)( ρV

ρL)

0.5

=( 38.56∗1825.71∗15 )∗( 71

1000 )0.5

=0.4796

From Fig 6.24(Seader), with a tray spacing of 24inches, CF = 0.23ft/s

Because FLV < 0.1, Ad/A = 0.1 and FHA = 1.0. Thus

C=(1.87 ) (0.75 ) (1 ) (0.23 )=0.32 fts

U f=0.32( 1000−7171 )

0.5

=1.16 fts

DT=( 4V MV

f U f π (1− Ad

A ) ρV )0.5

=(4∗25.72

3600 ∗15

0.80∗( 1.163.28 )∗3.14∗0.9∗71 )

0.5

=0.087m

24

PROCESS FLOW DIAGRAM

26

The figure above is the overall process flow diagram of the whole process. The FTR products from the reactor

are first sent to a stripping section where leftover methane, hydrogen and oxygen are stripped of the

hydrocarbon stream. The hydrocarbons are then sent to a settler where due to their density differences the water

can be separated off of the hydrocarbons. The hydrocarbons are then sent to a flash drum where tail gas is

separated and sent to the preheater as fuel. The bottom stream from the flash is then sent to another flash drum

where 50% LPG is recovered as gas and sold. The bottom stream of the 2nd flash is sent to a series of distillation

column where products are separated as ethane, LPG, and Naphtha. The bottom stream of the last distillation

column is then sent to hydro- isomerization unit where hydrogen is introduced and the heavier hydrocarbons are

broken into LPG, naphtha and diesel.

27

UNIT CONTROL AND INSTRUMENTATION

For Unit control and instrumentation a relief valve was sized for the Fischer Tropsch reactor for vapor service. Shown Below is the calculations and equations used for relief valve sizing.

In sizing reliefs in vapor service, the downstream pressure must be checked to ensure that it is less than the

choked pressure (Daniel, 2008).

Qm=Co∗A∗P √ ¿

Where

Qm is the discharge mass flow

Co is the discharge coefficient

A is the area of the discharge

P is the absolute upstream pressure

ɣ is the heat capacity ratio for the gas

gc is the gravitational constant

M is the molecular weight of the gas

Rg is the ideal gas constant

T is the absolute temperature of the discharge

The area of the relief vent given a specified mass flow rate Qm is;

A=Qm

C o∗P∗√ T

M/(( ɣ∗gc

Rg)∗( 2

ɣ+1 )ɣ+1ɣ−1)

Χ=√ ¿

The required relief vent area for an ideal gas is;

A=Qm

Co∗Χ∗P∗√ T

M

Χ=519.5∗√ ¿

The pressure, P is the maximum absolute relieving pressure. It is given for the

fixed-unit case by,

P=Pmax+14.7

From ASME Boiler and Pressure Vessel Code, the following guidelines are recommended for vapor reliefs.

28Pmax = 1.1Ps for unfired pressure vessels,Pmax = 1.2 Ps for vessels exposed to fire,Pmax = 1.33 Ps for piping

Using the input streams to the FT R, the area of the relief valve can be calculated.

From Chemcad simulation

Qm = 1169.75 lbm/h

T = 392 °F =

P = 362.59 psia

Cp/Cv =ɣ = 1.3683

M = 27.54 lbm/lb-mol

Χ=519.5∗√ ¿

Χ=519.5∗(1.3683∗( 21.3683+1 )

6.43)12 =352.92

Using the fire scenario for calculation

Pmax=1.2∗P s=1.2∗362.59=435.1 psia=P

Co = 0.975

A=Qm

C o∗Χ∗P∗√ T

M= 1169.75

0.975∗352.92∗435.1∗√ 851.67

27.54=0.0434 ¿2

D=( 4∗Aπ )

12=( 4∗0.0434 ¿2

π )12=0.2351∈¿

For our unit control we focused on the control of the FTR reactor as it is the most dangerous piece of equipment

in all of the plant. We came up with the following process control.

29

In the figure above there are many ways safety controls can be triggered. First of all the flow of boiler feed water is

controlled by temperatures from different parts of the reactor specifically the hot spots. From the calculations we found

out that the hotspots existed around the middle and at the end of the reactor. Next control is the liquid level of the utility

fluid. This is to ensure that there is no decrease in the liquid level and hotspot formation. The differential pressure sensors

also work in conjunction with the temperature sensor at the exit stream of the boiler feed water. The information from

these sensors are sent to an indicator alarm/DIC which then decides to close or open the incoming feed valve if there is

any spike in temperature or pressure difference. The closing of feed valve triggers the opening of a PSV in the feed stream

so as to direct the feed to an emergency flare. The flow of gas in the emergency flare then triggers the nitrogen purge into

the system. This flow of nitrogen then triggers the valve in the exit process stream to clos and PSV activates to direct the

contaminated process stream to a bypass vessel. This is done so as to avoid any contamination downstream.

SAFETY AND ENVIRONMENTAL SUMMARY

30

HAZARD ANALYSIS

The following table shows the potential hazards related to some chemicals in the process. The table is based on SAX’s dangerous properties of industrial materials.Materia

l

Health Hazards Fire/Explosion Hazard Corrosion/Reactivity Hazard

Methane

A simple asphyxiant High concentrations can cause nausea

and unconsciousness due to lack of oxygen

Very dangerous fire and explosion hazard when exposed to heat or flame.

Explosive in form of vapor when exposed to heat or flame.

Reacts violently with powerful oxidizers (e.g.: chlorine trifluoride, chlorine, fluorine)

Incompatible with halogens and interhalogens.

Carbon

monoxid

e

Toxic by inhalation can cause systemic changes like nausea, changes in blood pressure etc.

Prolonged exposure can cause heart disorders, nerve damage, reproductive effects, brain damage etc.

Blurred vision reported when in contact with eye

Very dangerous fire and explosion hazard when exposed to heat or flame.

Violent or explosive on contact with certain materials like iron oxide, chlorine dioxide, bromine trifluoride and liquid oxygen.

Can react with oxidizing materials, halogens, metal oxides, metals and lithium

Carbon

dioxide

Inhalation in large concentrations can cause rapid circulatory insufficiency and death.

An asphyxiant Contact with solid carbon dioxide

snow can cause burns

Minimal fire hazard (only reacts with certain materials like magnesium, zirconium, titanium etc and explodes).

Incompatible with acryladehyde, metal acetylides, sodium peroxide.

Hydroge

n

A simple asphyxiant Exposure to higher concentration may

lead to oxygen deficiency and may cause dizziness, drowsiness, nausea, loss of consciousness and death.

Highly flammable and explosive when exposed to heat, flame or oxidizers.

Flammable or explosive when mixed with air, oxygen and chlorine

Can react vigorously with oxidizing metals.

Forms explosive mixtures with bromine, chlorine, chlorine dioxide, liquid nitrogen etc.

Oxygen

Breathing high concentrations (< 75) can cause hyperoxia.

Liquid form can severely burn tissue due to extreme cold.

Noncombustible itself but essential to combustion

Can react with all flammable materials.

If proper conditions met can react with secondary alcohols, aluminum alloys, alkali metals.

C2-C4

Liquid form may cause frostbite An asphyxiant. Moderate

concentrations may cause headache, drowsiness and dizziness.

Lack of oxygen due to high concentrations of fuel gas may cause death.

Inhalation of butane may produce anesthetic effects and feeling of euphoria.

Highly flammable gas.(C2-C3) Forms explosive mixtures

with air and oxidizing agents. Gas may leak, spread and

create and explosive region. Flammable when combine

with air (C4).

The C2-C3 is normally stable. Avoid oxidizing agents,

chlorine dioxide, and chlorine.

31

Materia

l

Health Hazards Fire/Explosion Hazard Corrosion/Reactivity Hazard

Naphtha

Products

May cause lung damage if swallowed. Minor skin irritant can cause redness

or inflammation Inhalation may cause irritation to

respiratory system Mildly irritating to eyes

Highly Flammable Vapors can ignite rapidly

when exposed to ignition source.

Stable under normal conditions of use.

Avoid heat, flames and sparks and strong oxidizing agents

Diesel

Products

Contact with liquid or vapor may cause mild irritation

Causes skin irritation with prolonged contact

Aspiration may result in chemical pneumonia.

Can cause gastrointestinal disturbances, nausea and vomiting.

Can also effect the brain and severe cases can cause death.

Moderate fire hazard. Vapors can ignite rapidly

when exposed to ignition source.

Runoff to sewers may cause fire or explosion hazard

Mostly stable Incompatible with strong

oxidizers Can form hazardous

decomposition products.

32

The following table shows the safety and first aid measures related to the above mentioned hazards.

Material First Aid Measures Fire/Explosion Safety Environmental Safety

Methane Remove to fresh air if inhaled. If not breathing administer

artificial respiration and oxygen.

Wear a full face positive pressure self-contained breathing apparatus.

Use water spray to stop escaping gas and stop the flow of gas.

Follow proper procedures for waste disposal following federal, state and local environmental control.

Return in shipping container properly labeled to your supplier.

Carbon

monoxide

Remove to fresh air if inhaled If not breathing administer

artificial respiration and oxygen. If eye in contact with liquid

immediately flush eyes with plenty of water.

To fight fire use foam. Dry chemical and water to blanket fire.

Control containing vessels with water jets to reduce pressure and explosion

Do not dispose to sewers or water sources Follow proper procedures for disposal following

federal, state and local environmental control.

Carbon

dioxide

Remove to fresh air if inhaled If not breathing administer

artificial respiration and oxygen Flush with plenty of water if

exposed to skin or eye

N/A (noncombustible) Follow proper procedures for waste disposal following federal, state and local environmental control.


Do not dispose to sewers or water sources

Hydrogen

Remove to fresh air if inhaled Provide adequate oxygen if

exposed to high concentrations

Use dry chemical powders to control small fires.

Use water spray, fog or foam to control large fire.

Follow proper procedures for disposal following federal, state and local environmental control.

Oxygen Remove to fresh air if inhaled in high concentrations

Provide assisted respiration if unconscious.

Exclude itself from fire area of fire


Follow proper procedures for disposal following federal, state and local environmental control.

C2-C4

If skin or eye exposed to liquid immediately warm the frostbite area with warm water.

In inhaled immediately remove to fresh air.

Provide adequate oxygen if exposed to high concentrations

If digested do not induce vomiting and contact a physician

Immediately spray the fire with water from maximum distance.

Remove ignition source if without risk.

Evacuate all personnel from danger area.

May use Dry chemical extinguisher or water for butane fires



Do not dispose to sewers or water sources

33

Material First Aid Measures Fire/Explosion Safety Environmental Safety

Naphtha

Products

Remove the affected person to fresh air.

If not breathing apply artificial respiration.

Flush the contaminated skin with large amounts of water and use soap.

Flush eye with plenty of water if contaminated.

Isolate from heat source and naked flames.

Use foam, water spray or CO2 dry chemical powder for controlling fires

Recover and recycle if possible.

Contain spill with sand, earth or absorbent material.


Diesel

products

Immediately flush with clean water if exposed to eyes.

In inhaled immediately remove to fresh air.

Provide adequate oxygen if exposed to high concentrations

If digested do not induce vomiting and contact a physician immediately.

Remove contaminated clothing and wash the contaminated area with soap and water

Use any extinguisher suitable for small class B fires, dry chemical, water spray, firefighting foam etc.

Use water spray, fog or firefighting foam to fight large fires


Return in shipping container properly labeled to your

supplier. Do not dispose to sewers or water sources

Hazard Analysis for Unsafe Operation (HAZOP)

34

Equipment ProcessParameter

Guide Words

Deviation Causes Consequences Actions

Reactor

Level

More Less

Control valve for one of the incoming streams malfunctions

Clogged relief valveExplosions due to over pressurization, Process downtime

A pressure gauge should be installed between the spring relief valve and rupture disk

MoreLess

Pump FailurePump malfunctionNo Power

No control over feed flow to the reactor.Possible runaway reaction

Install flow indicator in the pipes leaving and entering the pump

Temperature MoreLess

Temperature of the reactor fluctuates

Too much or too less flow in utility stream.

Explosion of reactorRunaway reactionDisintegration of products

Install cascade of temperature and flow controllers in the reactor

Pressure MoreLess

Pressure inside the reactor deviates

Downstream compressor malfunctions, side reactions occurringTubes in reactor corrodes

Reduced conversion,Explosion hazardContamination of product

Install back up compressor, Install flow control valves and pressure controllers.

Distillation Column

Level More Less

Level indicator malfunctions or insufficient logging of feed data

Wear and tear in control valve.Lack of communication between operators

Adding too much feed to the process and causing over pressurization in the tower or adding too little and damaging the reboiler.

All indicators should be maintained and updated regularly and proper protocol should be followed when logging feed data


Temperature of the tower deviates from the set point

Failure of reboiler utility flow or failure of the condenser

Temperature in column reaches unsafe levels.Separation not achieved

Install control valves on both utility streams of condenser and reboiler

Pressure MoreLess

Overpressure in distillation column

Too much flow rate into the tower

Flooding the tower,Separation not achieved

Install control valves for the flow in and out of the unit op.

Heat Exchangers Flow

More Less

Supply too little or too much cooling/heating fluid

Wear and tear of control valves

Excess water is sent to the S/L separator

Install backup controllers

Equipment ProcessParameter

Guide Word

Deviation Causes Consequences Actions

35

s

Pumps Level MoreLess

Liquid level rises rapidly or decreases rapidly

Interlocks malfunctioning

Reduces the amount of feed produced.Too much flow or too less flow to unit op.

Making sure all interlocks are appropriately timed

Reflux Drums Level More

Less Tank LevelToo high or too flow rate to or from the reflux drums

Overflowing of tank and or over pressure.Cooling not achieved.

Relief valve should be installed, and an appropriate means of diverting excess fluid should be implemented

Absorbers

Level More Less Output flow rate

Wear and tear of control valve, or inappropriate dilution of corrosive compounds

Rapid corrosion of MOC, loss in HCl revenue

Care should be taken in calculating the required amount of water for dilution, and the appropriate MOC should be used


High liquid/gas flow temperatureLow liquid/ gas flow temperature

Temperature of the tower too high or too low

Low absorption.More/less utility flow rate.

Install temperature control indicator, set low and high alarms

Settler

Flow MoreLess

High or low flow rate into the system

Flow rate not controlledFlow control valves malfunction

Settling velocity is not met.Phase separation not achieved and product is contaminated

Install appropriate flow control valves upstream and downstream

Level MoreLess

High or low flow rate into the system

Flow rate not controlledFlow control valves malfunctionLevel controller malfunctions

Separation not achieved.Product may be contaminated by heavies or lights

Install appropriate level controllers in the settler.

36

ENVIRONMENTAL IMPACT

An environmental impact analysis was done on the waste water streams and the compounds in our process. By

utilizing the environmental impact tool on Chemcad we were able to evaluate each compound and its impact on

the environment. The tables showing the impacts are below:

Raw Component FactorsComponent Name Ozone Depletion Global

WarmingSmog Formation Acid Rain

Carbon Dioxide 0 1 0 0Methane 0 11 0.007 0

Water 0 0 0 0Hydrogen 0 0 0 0Oxygen 0 0 0 0Nitrogen 0 0 0 0

Carbon Monoxide 0 0 0 0Ethane 0 0 0.082 0

Propane 0 0 0.42 0n-Butane 0 0 0.41 0n-Pentane 0 0 0.408 0n-Hexane 0 0 0.421 0n-Heptane 0 0 0.529 0n-Octane 0 0 0.493 0n-Nonane 0 0 0.469 0n-Decane 0 0 0.464 0

n-Undecane 0 0 0.436 0n-Dodecane 0 0 0.412 0n-Tridecane 0 0 0 0

n-Tetradecane 0 0 0 0n-Pentadecane 0 0 0 0n-Hexadecane 0 0 0 0n-Heptadecane 0 0 0 0n-Octadecane 0 0 0 0n-Nonadecane 0 0 0 0

n-Eicosane 0 0 0 0

Raw Component FactorsComponent Name Human Toxicity Eco Toxicity

OSHA PEL LD 50 LC 50 LD 50

37

Carbon Dioxide 9000 0 24 0Methane 0 0 38.7995 0Water 0 0 0 0Hydrogen 0 0 0 0Oxygen 0 0 0 0Nitrogen 0 0 0 0Carbon Monoxide 55 0 15 0Ethane 0 0 95.7331 0Propane 1800 0 74.2984 0n-Butane 1900 1226.95 11.6925 1226.95n-Pentane 2950 1523.13 6.45919 1523.13n-Hexane 1800 28710 2.5 28710n-Heptane 2000 2115.28 1.96554 2115.28n-Octane 2350 2411.46 1.96735 2411.46n-Nonane 1050 2707.64 0.477075 2707.64n-Decane 0 3003.61 0.22227 3003.61n-Undecane 0 3300.01 0.167683 3300.01n-Dodecane 0 3595.76 0.330536 3595.76n-Tridecane 0 3892.16 0.593462 3892.16n-Tetradecane 0 4145.69 0.334588 4145.69n-Pentadecane 0 4484.31 0.194356 4484.31n-Hexadecane 0 4780.49 0 4780.49n-Heptadecane 0 5076.67 0 5076.67n-Octadecane 0 5372.64 0 5372.64n-Nonadecane 0 5668.82 0 5668.82n-Eicosane 0 5965 0 5965

Below, the environmental impacts of one of our waste water streams are shown:

38

Water Stream AnalysisComponent Name Flow

rate (kg/hr)

Ozone Depletion (impact/hr)

Global Warming (impact/hr)

Smog Formation (impact/hr)

Acid Rain (impact/hr)

Carbon Dioxide 0.01 0 4.76E-06 0 0Methane 0 0 5.41E-07 2.23E-06 0Water 187.54 0 0 0 0Carbon Monoxide 0.02 0 0 0 0

Water Stream AnalysisComponent Name Flow rate

(kg/hr)OSHA PEL (impact/hr)

LD50 (impact/hr)

LC50 (impact/hr)

LD50 (impact/hr)

Carbon Dioxide 0.01 4.85E-07 0 0.00135 0Methane 0 0 0.00 8.66E-06 0Water 187.54 0 0 0 0Carbon Monoxide 0.02 8.66E-05 0 2.36E-03 0

SAFETY CONSIDERATIONSIt is not enough to be able to design a plant that would be able to produce the desired products correctly; the

safety of the processes should also be put into consideration. In ensuring a safe process conditions, control

valves would be placed on all equipment to control pressure, temperature, and flow rate when needed.

Additional process modifications will also be carried out on important equipment. These modifications are

shown below;

Reactor

• Install a high-temperature alarm to alert the operator in the event of cooling function loss

• Install a high-temperature shut-down system .This system would automatically shut down the process in the event of a high reactor temperature

• Install a check valve in the cooling line to prevent reverse flow

• Study the cooling water source to consider possible contamination and interruption of supply

• Install a cooling water flow meter and low-flow alarm ( which will provide an immediate indication of cooling loss)

• Isolate Reactor with explosion walls

Drums

• Provide dip legs in all drums to prevent the free fall of solvent resulting in the generation and accumulation of static charge

All Equipment 39

• Ground all equipment

• Check the room periodically with colorimetric tubes to determine if any leaked vapors are present

• Increasing the ventilation with blowers

• Isolating process lines that contain hazardous streams

• Regular leak detection would be carried out

It is important to quantitatively carry out a risk analysis. The Chemical Process Quantitative Risk Analysis

(CQPRA) is a methodology designed to provide management with a tool to help evaluate process safety in the

chemical process industry. The steps1 involved are shown below;

In the event of leak from a vessel, the discharge rate can be quantified for liquid and gaseous flow.

From a general discharge rate model

Where

P= pressure (kPa)

ρ = density ( kg/m3)

g = acceleration due to gravity (m/s2)

gc = gravitational constant (N/kg-m/s2)

z = vertical height (m)40

v = fluid velocity (m/s)

f = frictional loss term

Ws = shaft work (J/s)

m = mass flow rate (kg/s)

ef = frictional loss term

Kf = excess head loss due to the pipe (dimensionless)

f = fanning friction factor

L = flow path length

D = flow path diameter

Using the 2k-Method to Estimate Kf

K f=k1

Nℜ+k∞∗(1+ 25.4

IDmin )Where, IDmin = internal diameter in mm

For pipe entrances and exits

K f=k1

Nℜ+k∞

For a normal entrance

k1 = 160

k∞ = 0.50

Border type entrance

k1 = 0

k∞ = 1.0

For high Reynolds number

NRe > 10000

Kf = k∞

For low Reynolds number

41

NRe < 50 ; Kf = k1/NRe

The fanning, f factor is calculated from the equation below;

1√ f

=−4 log( εD

3.7065−5.0452logA

N ℜ)

A={( εD )1.1098

2.8257+( 7.149

N ℜ)

0.8981}Where, ε = pipe roughness

For large Reynolds number

1√ f

=−4 log ( εD

3.7065 )For liquid discharge from a pipe assuming the flow is frictionless and that there is no shaft work would result in

a Bernoulli equation.

(P¿¿2−P1)1ρ+ ggc

∗( z2−z1 )+ 12∗gc

∗(v22−v1

2 )=0¿

The discharge of pure liquids through a sharp edged orifice

m=A∗CD∗√2 ρ∗gc∗(P1−P2 )Where,

m˙ = liquid discharge rate (kg/s)

A = area of hole (m2)

CD = discharge coefficient

ρ = density of fluid (kg/m3)

gc = gravitational constant (N/kg-m/s2)

P1 = pressure upstream of the hole (N/m2)

P2 = pressure of downstream of the hole (N/m2)

Tank Drainage Time

t= 1A∗C D∗√2g

∗∫V 2

V 1 dV (h )√h

Assumptions used in the equation above are;

Constant area

Constant discharge coefficient

Where,

42

t = time to drain tank from volume V2 to Volume V1 (s)

V = liquid volume in the tank above the tank (m3)

h = height of the liquid above the leakage (m)

Mass discharge rate of liquid from a hole in a tank

m=ρ∗v∗A=ρ∗A∗CD∗√2∗( gc∗Pg

ρ+g¿hL)

Where,

Pg = gauge pressure at the top of the tank (N/m2)

hL = height of liquid above the hole (m)

It should be noted that the mass discharge rate decreases with time as the liquid level drops. Therefore, the

maximum discharge rate happens when the leak first occurs.

Mass discharge rate of gas from a hole in a tank

Assumptions

Ideal gas

No heat transfer

No external shaft work

m=CD∗A∗P1∗√ 2∗gc

Rg∗M

T 1∗k

k−1∗{(P2

P1)

2k−( P2

P1)k−1k }

Where,

k = heat capacity ratio

Rg = ideal gas constant

T1 = initial upstream temperature of the gas

The mass discharge rate at choked flow is given as;

m=CD∗A∗P1∗√ k∗gc∗MRg∗T 1

∗( 2k+1 )(

k+1k−1 )

P choked

P1=( 2

k+1 )k

k +1

For ideal gas flow for both sonic and non-sonic conditions is represented by the Darcy formula and is given as;

m=Y∗A∗√ 2gc∗ρ1∗(P1−P2 )∑K f

Y = gas expansion factor (unitless)

43

Tempered Reaction

Q=m∗C v∗dT

dt+

(m∗V∗hfg )m∗v fg

Where,

Q = heat generation by reaction

m = mass within reactor vessel (kg)

Cv = heat capacity at constant volume (J/kg-°C)

T= absolute temperature of the reacting material (°C)

t = time (s)

hfg = enthalpy difference below liquid and vapor (J/kg-°C)

V = reactor vessel (m3)

vfg = specific volume difference between liquid and vapor

Heat release rate during overpressure

q=Cp

2∗{( dTdt ) , s+ dT

dt,m}

Where,

s = set conditions

m = turnaround conditions

For the quantitative analysis of the safety we considered a situation where there is a leak in oxygen power plant

and calculated the effects of the gas leak

Safety Calculation

Assuming that there is a leakage of gas in the Air separation unit, the discharge rate of the can be calculated to

determine if the leakage can be temporarily fixed or a shutdown of the system is needed.

The oxygen flow conditions are provided below;

P = 500 psig =34.77 bar

T = 75°F

Heat capacity = 1.37

P choked

P1=( 2

k+1 )k

k +1

P choked

P1=( 2

1.37+1 )1.37

1.37+ 1=0.9065

Pchoked=0.9065∗34.77=31.52 ¿44

Assume the diameter of the hole that is causing the leakage is 10mm

A=π∗D2

4=π∗(10mm )2

4=7.85∗10−5m2

( 2k+1 )

k +1k−1 =( 2

1.37+1 )1.37+11.37−1=0.3371

Assume CD = 0.85

mchoked=CD∗A∗P1∗√ k∗gc∗MRg∗T1

∗( 2k+1 )

k +1k−1

mchoked=.85∗7.85∗10−5m2∗34.77∗105 Pa∗√ 1.37∗ 31.96 kgkg−mole

8314 Pa∗m3

kg∗mole∗K∗296.89 K

∗0.3371

mchoked=0.5673kg

s

Initial flow rate of gas = 651040.8 kg/h

% lost=

1.5758∗10−4 kgh

651040.8 kgh

∗100 %=2.42∗10−8%

0.01 0.02 0.03 0.04 0.05 0.06 0.07 0.08 0.09 0.10

10

20

30

40

50

60

Mass flowrate Vs Diameter

Diameter(m)

disc

harg

e m

ass fl

owra

te(k

g/s)

As we can see from the figure the quantity of mass discharge can be quantifies. According to our calculations

we found out that at around 0.018 m diameter 1% of the mass is lost .This mass lost can be significantly high

taking in account the total amount of mass flow rate through the system.

45

HEAT AND MATERIAL BALANCE Mass balance was done according to kg/hr basis taking in consideration the molecular weight of each

component in the entire system. The following table summarizes the overall mass balance of the system.

46

Feed Kg/hr Syngas RKg/hr HE

Kg/hr Absorber 1

Kg/hr CO2 out Kg/hr

CO2 44.01 CO244.0

1 CO244.0

1 CO2 0.02 CO243.99

5

Water237.1

7 Water208.

7 Water21.1

6 Water 10.6 Water10.58

2

Oxygen393.5

9 Oxygen1.81

4 Oxygen1.81

4 Oxygen 0 Oxygen1.813

6

Nitrogen 3.446 Nitrogen3.44

6 Nitrogen3.44

6 Nitrogen 0 Nitrogen3.445

7

47

CO 0 CO730.

1 CO 730 CO 730 sum59.83

6Hydrogen

0 Hydrogen108.

3 Hydrogen108.

3 Hydrogen 108

Methane422.4

1 Methane4.22

4 Methane4.22

4 Methane 4.22

sum1100.

6 sum 1101 sum 913 sum 853

Out HEKg/hr

Water187.

5

FTR Kg/hr Absorber 2Kg/hr Top Out

kg/hr Settler

kg/hr Sett Out kg/hr

CO2 0 CO2 0 CO2 0 CO2 0 CO2 0

Water10.58

2 Water10.5

8 Water 0 Water 10.6 Water 10.58Oxygen 0 Oxygen 0 Oxygen 0 Oxygen 0 Oxygen 0

Nitrogen 0 Nitrogen 0 Nitrogen430.

8 Nitrogen 0 Nitrogen 0

CO430.7

6 CO430.

8 CO63.6

3 CO 0 CO 0

Hydrogen 63.63 Hydrogen63.6

3 Hydrogen37.4

6 Hydrogen 0 Hydrogen 0

C1 37.46 C137.4

6 C1 0 C1 0 C1 0

C2 0.388 C20.38

8 C20.01

6 C2 0.37 C2 0

C3 0.569 C30.56

9 C30.02

3 C3 0.55 C3 0

C4 0.75 C4 0.75 C40.03

1 C4 0.72 C4 0

C5 40.44 C540.4

4 C5 1.65 C5 38.8 C5 0

C6 44.91 C644.9

1 C6 1.84 C6 43.1 C6 0

C7 48.56 C748.5

6 C7 1.99 C7 46.6 C7 0

C8 51.47 C851.4

7 C8 2.11 C8 49.4 C8 0

C9 53.74 C953.7

4 C9 2.2 C9 51.5 C9 0

C10 55.43 C1055.4

3 C10 2.27 C10 53.2 C10 0

C11 56.63 C1156.6

3 C11 2.32 C11 54.3 C11 0

C12 57.38 C1257.3

8 C12 2.35 C12 55 C12 0

C13 57.75 C1357.7

5 C13 2.36 C13 55.4 C13 0

C14 57.78 C1457.7

8 C14 2.36 C14 55.4 C14 0

C15 57.53 C1557.5

3 C15 2.35 C15 55.2 C15 0

C16 57.02 C1657.0

2 C16 2.33 C16 54.7 C16 0

C17 56.31 C1756.3

1 C17 2.3 C17 54 C17 0

C18 55.41 C1855.4

1 C18 2.27 C18 53.1 C18 0

C19 54.36 C1954.3

6 C19 2.22 C19 52.1 C19 0

C20 53.19 C2053.1

9 C20 2.22 C20 51 C20 0

sum1402.

1sum

1402 sum567.

1 sum 835 sum 10.58

48

AsTo

49

s

Colmn 1 kg/hr Dist 1 kg/hr Dist 2 kg/hr Dist 3kg/hr

CO2 0 C20.35

6 CO2 0 CO2 0

Water 0 C3 0 Water 0 Water 0

Oxygen 0 C4 0 Oxygen 0 Oxygen 0

Nitrogen 0 C5 0 Nitrogen 0 Nitrogen 0

CO 0 C6 0 CO 0 CO 0Hydrogen 0 sum

0.356

Hydrogen 0 Hydrogen 0

C1 0 C1 0 C1 0

C2 0.37 Colm 2 kg/hr C20.01

7 C2 0

C3 0.55 CO2 0 C30.54

6 C3 0

C4 0.72 Water 0 C40.41

3 C4 0.31

C5 38.78 Oxygen 0 C50.00

1 C5 38.8

C6 43.08 Nitrogen 0 sum0.97

7 C6 43.1

C7 46.57 CO 0 C7 46.6

C8 49.37Hydrogen 0 Colm 3

kg/hrC8 49.4

C9 51.54 C1 0 CO2 0 C9 51.5

C10 53.17 C20.01

7 Water 0 C10 52.8

C11 54.31 C30.54

6 Oxygen 0 C11 2.33

C12 55.03 C40.71

9 Nitrogen 0 C12 0.01

C13 55.39 C538.7

8 CO 0 sum 285

C14 55.42 C643.0

7Hydrogen 0

C15 55.17 C746.5

7 C1 0 Botms 3kg/hr

C16 54.69 C849.3

7 C2 0 C10 0.36

C17 54 C951.5

4 C3 0 C11 52

C18 53.14 C1053.1

7 C40.30

7 C12 55

C19 52.14 C1154.3

1 C538.7

8 C13 55.4

C20 51.01 C1255.0

3 C643.0

7 C14 55.4

sum824.4

5 C1355.3

9 C746.5

7 C15 55.2

C1455.4

2 C849.3

6 C16 54.7

C1555.1

7 C951.5

4 C17 54

C1654.6

9 C1053.1

7 C18 53.1

C17 54 C1154.3

1 C19 52.1

C1853.1

4 C1255.0

3 C20 51

C1952.1

4 C1355.3

8 sum 538

C2051.0

1 C1455.4

1

sum824.

1 C1555.1

7

C1654.6

9C17 54

C1853.1

4 Kg/hr

C1952.1

3 Tot out 1203.9

C2051.0

2 Tot in1100.6

2sum 625

To HI

ENERGY BALANCE

The energy balance around the reactor was achieved by considering the enthalpy of all the incoming streams

and outgoing streams. The following table summarized the energy balance.

Overall Energy Balance MJ/h (Input) MJ/h (Output)

Feed streams -6.12E+06

Product -1.38E+07

Total Heating 2.05E+06

Total Cooling -1.00E+07

Power Added 5.54E+05

Power Generated -4.87E+04

Total -1.43E+07 -1.38E+07

From the table we calculated the % closure.

% closure=−1.38−1.43

∗100 %=96.51 %

CONCLUSION50

OPTIMIZATIONS AND RESULTS

The optimizations were done using the parametric optimization process. In the parametric optimization process

important decision variables were varied in order to obtain the most positive net present value in dollars. The

following variables were varied in order to achieve optimum conditions.

Feed preheat optimization

The feed preheat temperature entering the syngas reactor was optimized in this optimization. This was done

until an optimum value was achieved.

350 450 550 650 750 850 950 1050

-$3.85

-$3.80

-$3.75

-$3.70

-$3.65

-$3.60

-$3.55

NPV vs. Feed Preheat Temp

Preheat Temp. (ºF)

NPV

(Billi

ons o

f Doll

ars)

600 °F (-3.65 Billion)

From the optimization we found out that the optimum feeds preheat temperature was 600ºF at net present value

of -3.65 billion dollars.

Initial breakdown of NPV

51

The following pie chart shows the initial breakdown of COMd.

Cwt0%

Cut21%

Crm68%

Col0%FCI11%

Initial Breakdown of COMd

COM 4.29E+09

CWT 4.18E+06

CUT 9.10E+08

CRM 2.90E+09

COL 2.80E+06

FCI 4.70E+08

Revenue 4.72E+09

NPV -4.77E+09

From the initial breakdown of COMd we found out that the most expensive contributing factor was cost of raw

materials. As we had no control over the price of the raw material, we decided to optimize the consumption of

it. We also decide to optimize the FTR unit.

Syngas Unit Optimization

To ensure the optimum usage of raw material the syngas unit was optimized using different pressure

temperature and oxygen rate. This was done until optimum conditions were achieved.

52

1500 1550 1600 1650 1700 1750 1800 1850 1900 1950 2000

-5.00-4.00-3.00-2.00-1.000.001.002.003.004.00

NPV vs. SU (T/P/O2 rate)

Pressure 300 psig 2:1 molar ratio400 psig500 psig

Syngas Temp (ºF)

NP

V (

Bill

ions

of

$)

P (psig) T (°F) CO (kmol/h) H2(kmol/h) O2( kmol/h)

300 1600 15314 31806 27000

500 1950 20058 41076 20000

From the optimization we found out that the optimum condition was achieved at the pressure of 500 psig,

optimum temperature of 1950 F and oxygen flow rate of 20000 kmol/hr. The net present value obtained was

optimum at 2.63 billion dollars.

FTR unit optimization

The FTR unit reactor entering the FTR reactor was optimized in this optimization. This was done until optimum

conditions were achieved.

200 220 240 260 280 300 320 340 360

-3-2-1012345

FT Reactor Temperature

25 atm30 atm20 atm

Reactor Temp (⁰C)

NP

V (

$ B

illio

ns)

53

From the FTR unit optimization we found out that the optimum entering temperature into the FTR reactor was

251.7 ⁰C. This bumped out net present value to 4.2 billion dollars.

Final breakdown of NPV

The following pie chart shows the final breakdown of COMd.

Cwt0%

Cut35%

Crm49%

Col0%

FCI16%

Break down of COMd after SU/FT unit op-timization

COM 5.95E+09

CWT 7.64E+06

CUT 2.06E+09

CRM 2.36E+09

COL 2.80E+06

FCI 9.79E+08

Revenue 7.52E+09

NPV 4.95E+09

From syngas and FTR unit optimization as we can see the cost of raw material decreased but cost of utilities

increase. Nevertheless the NPV value also went up.

54

Effect if reactor trains in series

12

0

1

2

3

4

Effect of reactor trains in series

# of Reactors

NPV (

Billio

ns of

$) X= 30 %

X= 60 %

The above graph shows the effect of reactor trains. We increased the reactor train from just one to two in series

this made out conversion increase up to 60% and the optimum NPV obtained was 3.9 Billion dollars

Taking credit for utilities

Credit for utilitiesNo credit forutilities

3.84

4.24.44.64.8

5

Taking credit for utilities

NPV(

Billio

ns of

$)

We also added the credit for different utility streams and tail gas streams according to their heating values. This

drastically increased our NPV from 4.2 billion dollars to 4.96 billion dollars

55

Number of Trays/Feed Tray location optimization

The numbers of trays/feed tray locations were varied in the three distillation columns in order to achieve the

most positive net present value. This was done until an optimum was achieved.

4 5 6 7 8 9 10 11 12 13$3,863.60

$3,864.00

$3,864.40

$3,864.80

$3,865.20

$3,865.60

$3,866.00

$3,866.40

T-101 Optimization

151617182014

Feed Tray Location

NPV

($X1

06)

For T-101 a net present value of $3866.33X106 was obtained at 15 trays and feed tray location of 7.

4 5 6 7 8 9 10 11 12 13 14$3,891.00

$3,891.10

$3,891.20

$3,891.30

$3,891.40

$3,891.50

$3,891.60

T-102 Optimization

161718

Feed Tray Location

NPV

($X1

06)


56

4 6 8 10 12 14 16$3,891.56

$3,891.57

$3,891.58

T-103 Optimization

1920212223

Feed Tray Location

NPV

($X1

06)


Tie Spec optimization

After concluding on the working conditions for the separation process, we changed the recovery rate of the

various organics from each column. To achieve this tie spec for the reboiler of the distillation columns were

optimized by also varying the tie spec for the condenser.

0.88 0.9 0.92 0.94 0.96 0.98 1 1.023889.6

3890

3890.4

3890.8

3891.2

3891.6

3892

T-101:Tie Spec Optimization

Tie spec Reboiler 0.90.990.9990.9999

Tie spec for reboiler 1( MCB-TCB)

NPV

($X

106)

A maximum NPV value of $3891.554X106 was obtained at reboiler tie spec 0.99 mole fraction of and

condenser tie spec of 0.9999 mole fraction.

57

0.88 0.9 0.92 0.94 0.96 0.98 1 1.023890

3890.4

3890.8

3891.2

3891.6

3892


Tie spec Reboiler 0.999


NPV

($)

A maximum NPV value of $3891.554X106 was obtained at reboiler tie spec 0.999 mole fraction of and

condenser tie spec of 0.99 mole fraction.

0.2 0.4 0.6 0.8 1 1.2 1.4 1.6 1.8 23889.4

3889.6

3889.8

3890

3890.2

3890.4

3890.6

3890.8


Tie spec Reboiler 0.9

0.99

0.999


NPV

($)

A maximum NPV value of $3890.722X106 was obtained at reboiler tie spec 0.9 mole fraction of and condenser

tie spec of 0.3 reflux ratio.

58

ENERGY EFFICIENCY

We calculated the mass of carbon in the feed stream and exiting LPG, Naphtha, and Diesel streams by doing a

species balance on the carbon and estimating the closure with the following formula:

xkmol alkanehr

∗y kmolC

1kmolalkane∗12kgC

1kmolC

Hydrocarbon Carbon MassFeed CH4 2.99E+05LPGC3 6.90E+03C4 9.17E+03NaphthaC5 6.05E+03C6 6.97E+03C7 7.28E+03C8 2.43E+03C9 1.27E+03C10 6.85E+02DieselC11 2.06E+04C12 2.20E+04C13 2.29E+04C14 2.38E+04C15 2.45E+04C16 2.51E+04C17 2.56E+04C18 2.61E+04C19 2.65E+04C20 2.68E+04

Mass of carbon kgin 2.99E+05out 2.85E+05energy efficiency 95.23

The energy efficiency of our system is approximately 95%; we have assumed that the rest of the carbon is used up as fuel in the power generation unit or the unaccounted recycle stream..

59

RESULTS

We calculated the capital and operating cost of a GTL plant that utilizes a low temperature FT reactor. After performing all the optimizations we found out that these are the best conditions for our GTL plant. The following table summarizes the best case conditions:

Economics Summary

Best case $ Cost of Raw materials ($/yr.) 2,360,000,000

NPV 4,960,000,000 Methane 339,000,000

FCI 5,200,000,000 Oxygen 543,000,000

Top 5 equipment (name/cost) Carbon Dioxide 9,300,000

1 FTR 33,207,000 Steam 7,900,000

2 E-109 2,800,000 Catalyst 40,260,000

3 E-105 1,560,000 MEA 879,000,000

4 T-102 900,000 Waste treatment ($/yr.) 6,220,000

5 T-103 832,000 Revenue Total ($/yr.) 7,520,000,000

COM ($/yr.) 5,903,400,000 LPG 526,000,000

Top 5 utilities (name/dollars) Naphtha 1,410,000,000

1 FTR 290,000,000 Diesel 5,589,000,000

2 Oxygen

Plant

217,000,000 Credit For Utilities ($/yr) 18,287,000

3 E-101 183,000,000 1 Fuel gas 1,370,000

4 E-109 61,000,000 2 lps 4,800,000

5 E-105 60,300,000 3 mps 11,800,000

4 Steam Condensate 407,000

Operating Labor ($/yr) 1,240,000 Thermodynamics MSRK

60

After performing all the optimizations and calculation our final NPV value obtained was $ 4.96 Billion. The pie chart below shows different utility breakdown; which is followed by equipment specifics of final design.

E-10116% E-102

3%

E-1037%

E-1040%

E-1055% E-106

1%E-107

2%E-108

0%E-109

5%

R-10126%

R-10215%

H-1011%

C-1010%

P-101 A/B0%

P-102 A/B0%

P-103 A/B0%

C-1020%

C-SU10%

O2 plant19%

Utility Breakdown

The most contributing utility in our final design were the two reactors followed by oxygen plant and syngas cooling heat exchanger train.

NPV $4.96E+9

Feed Ratio (Hydrogen/Carbon

monoxide) 2:1

Reactor 1

Volume 533 m3

Pressure 25 atm

Temperature 500 of

Reactor 2

Volume 533 m3

Tower 101

Number of Trays 15

Feed tray location 7

Tower 102

Number of Trays 16


Tower 103

Number of Trays 22


Tie spec for T-101

Distillate Comp mole fraction 0.9999

Bottoms Comp mole fraction 0.99

61

Tie spec for T-102

Distillate Comp mole fraction 0.99


Tie spec for T-103

Distillate Reflux Ratio 0.3


RECOMMENDATIONS

We used a distillation column to separate the tail gas (ethane) from the rest of the FT products. This method

proved cumbersome, because we either had to operate the column at pressures above 300 psia, or use a

refrigerant at -50ºC. Both alternatives were expensive; we would recommend that a different means of

removing the ethane be researched. Another alternative to our process will be to operate the FT reactor at

temperatures that favor the product that is in demand. For instance, the FT reactor produces more diesel at

temperatures between 200 and 250ºC and more of the lighter alkanes at higher temperatures. Also we have

evidence that suggests the separation of LPG from Naphtha, and Diesel can be done in a fractional distillation

column. The product selling price in the problem statement is not current. The increase of price of the product

to current values will positively impact the NPV. However as the fuel prices increase there will inevitably

increase in utility and raw materials but from the following graph we see that 25% increase in selling price will

increase the NPV by 79%.This will definitely impact NPV positively even with the increase in utility and raw

materials.

0.30.4

0.00 5.00

10.00 15.00 20.00 25.00

25% increase in Product Selling Price

Selling Price ($/lb)

NP

V (

Bill

ions

of

$)

Naphtha $100/bblDiesel $118/bbl 79% Increase in NPV (18.82 Billion)

62

Also investigating production of oxygen with the plant is another option to consider. As oxygen was the most

expensive utility production of oxygen will drastically reduce raw materials cost. Also there are different

methods of syngas production which are more efficient. From our research we found out one of them is auto

thermal reforming .This can also impact NPV. Another option is investigating catalyst life. From a quantitative

analysis we found out that doubling the catalyst life increase the NPV by 6%.The graph below shows the

relation.

4 8

4.50

5.00

5.50 Effect of Doubling Catalyst Life

Catalyst Life

NP

V(B

illio

ns o

f $)

6% increase in NPV (317 million)

63

REFERENCES1. NACE International, The Corrosion Society (2002). Corrosion Survey Database (COR•SUR)..

NACE International. Online version available at:

http://www.knovel.com/web/portal/browse/display?

_EXT_KNOVEL_DISPLAY_bookid=532&VerticalID=0

2. Crowl, Daniel A., and Joseph F. Louvar. ChemicalProcessSafety:Fundamentalswith

Applications. Upper Saddle River, NJ: Prentice Hall PTR, 2008. Print.

3. Seader, J. D., and Ernest J. Henley. SeparationProcessPrinciples. Hoboken, NJ: Wiley, 2006. Print.

4. Smith, R. ChemicalProcessDesign. New York: McGraw-Hill, 1995. Print.

5. McCabe, Warren L., Julian C. Smith, and Peter Harriott. UnitOperationsofChemical

Engineering. Boston: McGraw-Hill, 2005. Print.

6. Turton, Richard. Analysis,Synthesis,andDesignofChemicalProcesses. Upper Saddle

River, NJ: Prentice Hall, 2009. Print.

7. Fogler, H. Scott. ElementsofChemicalReactionEngineering. Upper Saddle River, NJ:

Prentice Hall PTR, 2006. Print.

8. Swanson, Ryan M., Justinus A. Satrio, and Robert C. Brown. Techno-EconomicAnalysis

ofBiofuelsProductionBasedonGasification. Tech. U.S. Department of Energy. Web. 8

Mar. 2011. <http://www.nrel.gov/docs/fy11osti/46587.pdf>.

9. BaselineDesign/EconomicsForAdvancedFischer-TropschTechnology.Quarterly

Report. Tech. U.S. Department of Commerce. Web. 15 Mar. 2011. <http://www.fischer-

tropsch.org/DOE/DOE_reports/90027/90027_01/90027_01_toc.htm>.

10.Fixed-BedCatalyticReactors. Nob Hill Publishing, LLC.

11.Beiermann, Dipl.-Ing. Dagmar. FromFischerTropschRawproductstoFischerTropsch

Fuels:DevelopmentofanUpgradingModelandApplicationtoXtLProcesses. Rep.

DGMK. Web. 15 Mar. 2011.

<http://www.dgmk.de/petrochemistry/abstracts_content15/Beiermann.pdf>.

65

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