HO CHI MINH CITY UNIVERSITY OF TECHNOLOGY

OFFICE FOR INTERNATIONAL STUDY PROGRAMS

——————– * ———————

PROJECT REPORT
CHEMICAL PROCESS DESIGN – CH3321
SEMESTER: HK242

DRYING OIL PRODUCTION

LECTURER: NGUYỄN THÀNH DUY QUANG

CLASS: CC02 – GROUP: 3

FULL NAME STUDENT’S ID

Nguyễn Ái Minh Châu 2252089
Nguyễn Phạm Quốc Duy 2153253
Nguyễn Đức Hiếu 2252210
Nguyễn Ngọc Như Quỳnh 2252704
Lý Vĩnh Anh Quân 2252677
Trương Hữu Tài 2252729

HO CHI MINH CITY, MAY/2025

 OUTLINE
I. EXECUTIVE SUMMARY...............................................................................1
II. INTRODUCTION.............................................................................................3
III. TECHNICAL APPROACH.............................................................................4
IV. DESIGN SUMMARY........................................................................................5
1. Overview.............................................................................................................5
2. HYSYS Simulation.............................................................................................5
3. Pre-reaction process...........................................................................................6
4. Reaction stage.....................................................................................................6
5. Separation...........................................................................................................6
6. Recycling ACO...................................................................................................7
7. Purification.........................................................................................................8
8. Process Optimization..........................................................................................8
V. PROCESS ALTERNATIVES..........................................................................9
1. Reaction process.................................................................................................9
2. Separation.........................................................................................................10
3. Purification and recycling...............................................................................11
VI. ECONOMIC ANALYSIS...............................................................................13
1. Capital Costs.....................................................................................................13
2. Equipment Cost Distribution...........................................................................13
3. Energy Usage....................................................................................................14
4. Annual Operating Cost - AOC.........................................................................15
5. Overall Economic Feasibility...........................................................................16
6. Net Present Value (NPV).................................................................................17
VII. CONCLUSION AND RECOMMENDATION.............................................18
VIII. REFERENCES................................................................................................19
IX. APPENDICES..................................................................................................20
1. Appendix: Process flow diagram.....................................................................20
2. Appendix: HYSYS diagram.............................................................................20
3. Equipment list...................................................................................................21
4. Economics analysis..........................................................................................22
 LIST OF TABLES
Table 1: Distillation Columns Data.............................................................................21
Table 2: Separator Data..............................................................................................21
Table 3: Reactor Data.................................................................................................21
Table 4: Heaters, Condensers and Reboilers Data.....................................................21
Table 5: Power of pumps.............................................................................................22
Table 6: Energy Cost Table.........................................................................................22
Table 7: By Energy Type..............................................................................................23
Table 8: Equipment Scaling Data................................................................................24
Table 9: Cost by Equipment Type................................................................................24
Table 10: The project pays off quickly after just over 1 year of operation..................24

LIST OF FIGURES
Figure 1: Process flow diagram of Drying Oil Production.........................................20
Figure 2: HYSYS diagram of Drying Oil Production..................................................20
Figure 3,4: Preliminary design and information of the flows through the distillation
column T-100................................................................................................................25
Figure 5,6: Preliminary design and information of the flows through the distillation
column T-101................................................................................................................26
I. EXECUTIVE SUMMARY
This report presents the comprehensive design and simulation of a drying oil
(DO) production system - an industrially significant product known for its self-drying
properties. With moderate viscosity and excellent film-forming ability, drying oil
finds widespread application as a binder in paints, inks, and varnishes. The process
utilizes acetylated castor oil (ACO) as the primary feedstock. Due to the complex
chemical nature of both ACO and DO, simplified surrogate compounds were used for
simulation purposes: ACO is represented by palmitic acid (C 15H31COOH), and DO is
modeled as 1-tetradecene (C14H28). An undesirable side reaction can lead to the
formation of a polymeric by-product - modeled as 1-octadecene (C28H56) commonly
referred to as "gum" because of its sticky properties. The feed stream entering the
system contains approximately 1621 kg/h of pure ACO. The process generates two
marketable products: drying oil and acetic acid (AC), the latter being a commercially
valuable by-product.
The system design is divided into four major sections, all simulated using Aspen
HYSYS: reaction, separation, recycling and purification. Initially, a preparation stage
blends the fresh ACO feed with recycled ACO and stabilizes in a tank. This mixture is
then preheated and introduced into the reactor, where the DO is synthesized, along
with by-product Acetic Acid and the side reaction product - gum. The final conversion
of ACO is approximately 17.83%. The resulting two-phase stream - comprising solid-
phase microplastics dispersed in a liquid matrix - is then being cooled. Following the
reaction, the separation stage begins. The solid gum particles are removed through
filtration, and the liquid phase proceeds to a distillation column to recover unreacted
ACO from the DO-Acetic Acid mixture. Due to the differing boiling points of these
components, indirect distillation is used for effective separation. The bottom stream of
the first column, consisting of unreacted ACO, is recycled back into the system,
optimizing raw material use and minimizing costs. The overhead stream is fed into a
second distillation column, where DO is taken from the bottom, and acetic acid is
collected at the top.
This discussed design results in the following product streams: 1227 kg/h of
99% drying oil, 392.1 kg/h of high-purity 99% Acetic Acid. The estimated annual

1
revenue from product sales is approximately 567.63 million USD. The capital
investment in equipment is estimated at 10.10 million USD, while annual operating
costs amount to 526.99 million USD. With a payback period of only 1.18 years and a
profitability index (PI) of 5.21, the system demonstrates high financial viability and
attractive returns for investors. However, as the results are based on simulation,
further research and pilot-scale development are recommended to ensure practical
feasibility and minimize risks.

2
II. INTRODUCTION
This report summarizes a project to design a process to produce ethylene and
NGLs from shale gas. The production of Drying oils is desirable because Drying oils
play a critical role in the production of industrial coatings such as paints and
varnishes. Their unique self-hardening characteristic is due to oxidative crosslinking,
which forms a durable film upon exposure to air. In addition to the primary product,
the process also generates acetic acid, which can be commercialized, thus improving
the overall economic efficiency of the system. This study employs Aspen HYSYS to
design and analyze the drying oil production process. Through simulation, optimal
process parameters and design choices were identified to maximize efficiency, reduce
capital investment, and minimize operational costs - ultimately yielding a cost-
effective, high-output production system. In addition to technical design, this report
provides a thorough economic analysis, including assessments of capital investment,
fixed and variable operating costs, and net present value (NPV). A sensitivity analysis
was also conducted to evaluate the impact of alternative design choices and market
fluctuations on the economic performance of the project.

3
III. TECHNICAL APPROACH
The feedstock for the process is castor oil, extracted from castor seeds through
mechanical pressing and subsequently acetylated. The total input flow of ACO into
the system is 1621 kg/h, and after mixing the recycle stream, the total flow rate before
entering the reactor is 9147 kg/h with a trace amount of DO. The production process
does not rely on a single chemical reaction but rather involves a series of thermo-
mechanical operations to isolate and concentrate the desired components. The mixture
is subjected to indirect distillation in two columns - T-100 and T-101 - to efficiently
separate volatile compounds and unreacted feedstock from the final products. The
outcome is a two-product system: 99% pure drying oil, suitable for industrial use,
99% pure acetic acid, suitable for commercialization. Additionally, a minor amount of
waste product (referred to as "gum") forms as a result of unwanted polymerization.
This by-product is effectively removed through filtration to maintain the purity of the
final output. The simulation results from HYSYS demonstrate that the process is
technically viable and economically attractive. The optimized process design ensures
minimal waste, high product yields, and low operational overhead, thereby offering a
strong return on investment for stakeholders.

4
IV. DESIGN SUMMARY
1. Overview
The proposed design consists of 4 main sections: Reaction, Seperation,
Recycling and Purification. ACO is initially fed from a holding tank, where it is
blended with recycled ACO to form a combined feed. This mixture is then heated to
the required reaction temperature in the fired heater. The reaction proceeds without
the need for a catalyst and takes place in the reactor R-100, which is a vessel equipped
with inert packing to promote effective radial mixing. Upon exiting the reactor, the
reaction mixture is rapidly cooled (quenched). The mixture is then sent to the
separator F-100, where the solid small particles are separated from the liquid phase by
filtration. The filtered liquid is pumped to a distillation column for component
separation. In the T-100 distillation column, ACO is separated from the mixture of
DO and acetic acid. The bottom product containing mostly unreacted ACO and the
light end DO will be recycled and mixed with the inlet pure ACO stream, the top
product will be transferred into the second column for separate ACO and Acetic Acid.

2. HYSYS Simulation
The design described above was simulated in Aspen HYSYS to determine
necessary stream specifications, required utilities and sizing information for the
various pieces of equipment for this process. A picture of the setup in HYSYS and
details of the HYSYS results, including stream and equipment specifications, can be
found in the Process Flow Diagram.
The fluid package used for the simulation was Peng-Robinson Stryjek-Vera
(PRSV), an extension from PR which was chosen because it is most compatible with
gas components. In practical applications, PRSV is commonly used for applications
such as natural gas processing, gas separation, and processes involving both
hydrocarbon and non-hydrocarbon substances. It is particularly effective in low-
temperature applications, where accurate phase equilibrium calculations are crucial.

5
 3. Pre-reaction process
In this stage, the feedstock is prepared for the reaction. A pure Acetylated Castor
Oil (ACO) feed stream (1621 kg/h) is thoroughly mixed with recycled ACO and
stabilized in the storage vessel. The resulting mixture is then pumped into the system.
This stage operates at 25°C and 110 kPa. The combined feed stream has a total flow
rate of 9147 kg/h, consisting mainly of ACO (99% mole fraction) with a small amount
of DO from the recycle stream. This mixed feed is then pumped through a heater,
raising its temperature from 147°C to 380°C before it enters reactor R-100.

4. Reaction stage
The heated ACO mixture is fed into reactor R-100. Here, ACO is converted into
DO as the primary reaction, alongside a side reaction that produces DO
polymerization. The reactor R-100 outlet stream is a liquid containing a small amount
of dispersed fine solid particles (gum), a byproduct of the side reaction. This stream
contains unreacted ACO (69.26% mole fraction), the main product DO (15.71% mole
fraction), and the byproduct Acetic Acid (15.03% mole fraction). The exothermic
reaction causes the product stream temperature to drop to 342.8°C. The composition
of the outflow shows a significant increase in DO, with some unreacted ACO
remaining. Acetic Acid is also generated as a byproduct. The ACO conversion rate
reaches 17.83%. This conversion value is designed to optimize costs, considering
residual ACO reflux, ACO input, and reactor fabrication costs. Finally, the reactor
outlet stream is pumped into a cooler, reducing its temperature from 334.4°C to
180°C.

5. Separation
After being cooled, the components are transported at a temperature of 180°C
and a pressure of 148 kPa to two holding vessels. Two vessels are connected in series.
The post-reaction stream, which is a heterogeneous mixture of solid and liquid phases,
first enters Vessel 1 (not shown), which functions as a filtration unit to separate solids
from the liquid products. It is connected to the gum storage unit via stream line 8, with
a flow rate of 1.818 kg/h. The clarified liquid then flows into Vessel 2 (X-100), which

6
acts as a temporary storage or holdup tank. From there, the liquid (with composition in
mole fraction of 0.69 ACO, 0.16 DO and 0.15 Acetic acid) is continuously pumped to
the distillation column at 9145 kg/h for further separation.
The main purpose of X-100 is to provide buffer storage of the filtered liquid,
ensuring continuous downstream operation in the event that Vessel becomes clogged
(e.g., due to gum or solid accumulation). This allows maintenance tasks such as filter
screen replacement to be carried out without interrupting the entire process.

6. Recycling ACO
The separation system is designed to separate a mixture of three components:
acetylated castor oil (ACO), acetic acid, and drying oil (DO) using an indirect
sequence of distillation. The system has a series of 2 distillation columns, while the
first one T-100 is used to recycle unreacted ACO. T-100 has 10 trays, while the inlet
is fed to tray number 5 and it is operated at 122.2 kPa, higher than 260oC for the
upper position as well as at 136.0 kPa, about 200oC for the lower from tray 5. The top
product is a vapour combination of Acetic acid and DO. This stream is then sent to
Cond.1 in which the stream is condensed and then becomes liquid again. After that,
components are kept in a recycle tower reflux drum before being pumped back again
to the column with reflux ratio being 1. Then, the distillate goes to the next distillation
column at 1619 kg/h mass flow rate, with the composition of 50% of DO and 50% of
Acetic acid in mole. The bottom product is a liquid stream of ACO after which is sent
to recycle tower reboiler Reb.1 at 348oC, 90kPa. This stream with a mole fraction of
0.99 ACO is then sent into the next recycling step at 7526 kg/h.
The next step is to cool down the ACO stream in cooler E-102, which reduces
the temperature from 362.6oC to 170oC and reduces the pressure from 136 kPa to
135.5 kPa. After that, the fluid stream is pumped by the pump P-101 to the storage
tank for recycled ACO RCY-1 which is operated at 170 oC and 144.8 kPa. Then, it
containing 99% mole of ACO will be supplied to the MIX-101 with a flow rate of
7527 kg/h, where it mixes with the pure ACO feed stream.

7
 7. Purification
The second distillation column, T-101, separates drying oil from the mixture of
drying oil and acetic acid, which come from the top of the column T-100. The fluid
stream, which has the mole fraction of 49.96 % drying oil and 50.04% acetic acid,
comes in at a rate of 1619 kg/h, both with a temperature around 133.5 oC and pressure
of 122.2 kPa. The column operates at pressure varies from 108.5 kPa and 122.2 kPA
and has 10 trays, with the feed entering on the fifth tray from the top. After
distillation, the top product stream, which contains mainly acetic acid of 99%, exits at
392.1 kg/h at a temperature of 119.7 oC and pressure of 108.5 kPa. While the drying
oil, the desired product, is taken off the bottom of the tower and has a purity of
approximately 100 wt%, it comes out at a rate of 1227 kg/h, with a temperature of
260.8oC and pressure of 122.2 kPa.

8. Process Optimization
The process was optimized to minimize total cost, encompassing both capital
and operating expenses. Specifically, the reflux ratio and the number of trays in each
distillation column were optimized to reduce the initial capital investment while
meeting product purity specifications. Utility usage, a significant operating expense,
was minimized by designing each column to operate at the highest possible pressure,
which reduces energy required for condenser and reboiler duties.
To further optimize overall costs, the process incorporates a recycled stream of
unreacted ACO. This approach allows for a comprehensive evaluation of equipment
fabrication, raw material input, and ACO reflux costs to ensure the lowest possible
total production cost. Additionally, a liquid product storage tank is employed after
filtration to prevent process interruptions during filter section maintenance in the
filtration tower. This tank provides a buffer, allowing the process to continue running
smoothly during filtration system maintenance.

8
V. PROCESS ALTERNATIVES
1. Reaction process
Multi-stage reaction
A potential alternative to the current reaction process is the application of a
multi-stage reaction model. Instead of using a single reactor, the process would be
divided into two or more reactors connected in series. After leaving Reactor 1, the
unreacted product stream would be sent to Reactor 2 for further conversion, with
optimal operating conditions specific to each stage, such as temperature, pressure, and
flow rate. The main objective of this approach is to increase the conversion of ACO to
DO, while also better controlling the side reactions that lead to by-products such as
acetic acid and gum. A key advantage of the multi-stage reaction is the ability to
increase overall conversion, from the current 17.83% to a higher value, by
sequentially processing the unreacted material. Dividing the process into multiple
stages also helps to better control side reactions and reduce the formation of
undesirable products. Additionally, the heat generated by the first reactor can be
utilized to provide heat to the subsequent reactor, contributing to energy consumption
optimization. However, this approach also comes with certain disadvantages. First, the
initial investment cost increases, the system becomes more complex, and it requires a
larger plant area to accommodate the series of reactors. Therefore, this method was
not selected.

Atomization techniques
Another feasible approach to accelerate the formation of DO is to apply
atomization techniques. If the ACO feed is introduced into the reactor as a mist or
microjet stream, the surface area for contact increases significantly, promoting both
heat and mass transfer. The main advantage of this method is that it can substantially
enhance the reaction rate and improve the conversion rate without requiring a
significant increase in operating temperature. Moreover, temperature distribution
within the reactor becomes more uniform, reducing the risk of hot spots and
minimizing the formation of byproducts. However, the downside is that the system
becomes more complex, requiring the addition of specialized nozzles and strict control

9
over droplet size and spray flow rate. Additionally, the risk of clogging due to gum
formation during operation must be carefully considered and properly managed.

2. Separation
Hydrocyclone Gum Separation

A hydrocyclone can utilize the gum separation step due to centrifugal force,
which is better to deal with abrasive or sticky solids like gum. The feed mixture enters
the hydrocyclone tangentially at high velocity, creating a spiral (vortex) motion.
While heavier solid particles (like gum) are forced to the outer walls and downward
due to centrifugal force, exiting through the underflow, lighter liquid components
move to the center and exit through the overflow at the top.
This equipment can make the process become continuous without interruption
by replacement or cleaning, as well as reduces clogging issues common with sticky
gum in traditional filters. In addition, the energy supplied and operating cost might be
lower. On the other hand, the hydrocyclone is less effective for extremely fine
colloidal form of gums, which can affect the further separation efficiency, therefore it
is not practical to use in this process

Catalytic Reactor to Minimize Gum Formation

Gum, an undesired product, with less value in the economic path should be
minimized while the current process is let it be produced then be separated. There are
several ways that minimizing gum formation can be accomplished. One of those is the
replacement of the current reactor CVR-100 by a catalytic one. Gum formation can be
reduced by adding a solid acid catalyst, such as zeolite or silica-alumina, to the reactor
in place of the existing inert packing. This will lower the reaction temperature and
suppress the side reaction. By lowering solid waste, this improves DO output, lowers
energy use, minimizes gum production and simplifies downstream filtration.
However, due to the fact that the most challenging step here is finding the appropriate
catalysts with the suitable amount and properties. Catalyst development and testing
are necessary to guarantee selectivity and stability because of the possibility of
deactivation, the need for frequent replacement and the possibility of catalyst

10
poisoning by impurities, which could reduce reaction efficiency and indicate for strict
monitoring and maintenance.

3. Purification and recycling

Spray Drying for Acetic Acid Removal
To simplify acetic acid separation, the second distillation tower can be replaced
with a spray drying unit that atomizes the overhead stream from the first distillation
tower into a drying chamber, where hot air evaporates acetic acid ,which has a boiling
point 118°C, leaving DO, which has a boiling point ~250°C, as a powder or liquid for
storage. This eliminates the second tower’s high energy demands, simplifies the
process by removing the tower and reduces maintenance needs due to fewer moving
parts. However, acetic acid vapor recovery via a condenser or scrubber adds
complexity, the spray dryer may alter DO’s physical form as a powder or liquid,
potentially impacting its use in paints, and the high operating temperature risks
thermal degradation of DO if not controlled precisely.

Vacuum distillation for ACO recovery

Instead of simple distillation to separate unreacted ACO from DO and Acetic
acid, another option would be vacuum distillation. Under vacuum, relative volatility
differences improve, making separation more efficient with higher purity. Therefore,
the system needs to be operated at lower pressure leading to lower energy
consumption for heating at very high temperature. However, the operational
complexity and high cost to operate the vacuum system make it not used.

Single Distillation Tower with Side Draw

Instead of using two distillation towers to separate DO, acetic acid and ACO,
another option would be to combine those towers into a single tower with a side draw
at tray 25 for DO can streamline separation while recovering ACO at the bottoms and
acetic acid at the overhead. A single tower with 60–70 trays and a side draw at an
intermediate tray could achieve similar separation of those two. Besides reducing
equipment count, minimizing maintenance needs and decreasing utility demands by

11
using less steam and cooling water, this also reduces the plant’s footprint as well as
simplifying layout. However, there are some challenges that could be considered such
as the complex design and control to achieve three-product separation in one tower, a
higher reflux ratio may raise energy usage and there’s a risk of reduced DO purity if
tray optimization fails, requiring advanced control systems and potentially affecting
product quality.

12
VI. ECONOMIC ANALYSIS
1. Capital Costs
In the energy cost section, our team proposed combining two estimation
methods: Scaling and the Lang Factor method. The capital cost is primarily
determined by the Purchased Equipment Cost (PEC), which is then multiplied by a
Lang Factor to estimate the Total Capital Investment (TCI). The Scaling method
further refines the accuracy and specificity of TCI estimation based on Lang factors:
TCI =F L × ∑ PEC i
Where the Lang Factor typically ranges from 3.1 to 4.7 for continuous chemical
plants. Since Drying Oil Production is a continuous liquid processing operation, it
falls under the Fluid Processing category. We therefore selected a Lang Factor of 4.7.
When equipment capacity information is available, the Scaling Method can be
applied:
C=C 0 ¿
Where:
● C 0 is the known cost at base capacity S0
● S is the design capacity
● n is the scaling exponent (typically 0.6–0.7 for mechanical equipment)
The capital investment cost was estimated using this scaling formula combined
with the Lang Factor. The total PEC amounts to 10,104,045.73 USD, and applying the
Lang Factor of 4.74 results in a Total Installed Cost (TIC) of 47,893,176.74 USD. The
ratio PEC/TIC = 21.10% falls within the typical range for the chemical industry (15%
– 25%), affirming the technical reliability and economic feasibility of this estimate.
This validates the appropriateness of the adopted methods and assumptions for
continuous liquid processing plants.

2. Equipment Cost Distribution

Distillation Column accounts for the largest share at 82% of PEC (~5,249,810
USD), reflecting its central role in drying oil processes where separation and
distillation govern product quality and process efficiency.

13
 Shell & Tube Heat Exchangers make up 17.2% of PEC (~1,098,253 USD),
corresponding to the high heat exchange demand.
Centrifugal Pumps and Vertical Vessels represent minor shares at 0.1% and
0.8%, indicating a process reliant on thermal separation rather than power-driven or
pressurized operations.
This cost distribution aligns with the continuous chemical processing nature and
allows designers to focus on optimizing high-investment units for maximum return.

3. Energy Usage
The total energy cost is commonly expressed as:
n
Energy cost =∑ Qi × Ei
i=1

Where:
● Qiis the heat or electrical load at each equipment unit (GJ/h or kWh)
● Ei is the unit energy price (USD/GJ or USD/kWh), often obtained from Aspen
simulation.
The annual operating energy cost (over 20 years) of the project was calculated
based on the actual total consumption capacity of each equipment group,
corresponding unit energy price, and the assumption that the plant operates
continuously for 8766 hours/year.
The results show a total energy cost of 1,190,945 USD/year, in which the largest
share is for the Cooling Water group with over 1,101,408 USD/year, equivalent to
~92.5% of the total energy cost. This accurately reflects the operational nature of the
Drying Oil line, where the heat exchange and cooling systems operate at high
frequency and intensity to maintain temperature stability throughout the continuous
drying and distillation process.
Fuel/Steam costs account for 88,516 USD/year, equivalent to 7.4%, indicating
that although the process requires high thermal energy, the heating stages have been
mostly optimized through internal heat exchange and energy recovery, reducing new
fuel consumption. Meanwhile, Electricity costs are the lowest, at about 21 USD/year,
almost negligible compared to the other two categories. This structure indicates that

14
the project has a low electricity demand, mainly for pumps and auxiliary devices,
while thermal energy is the primary deciding factor. This particularly suits the
technological characteristics of a continuous Drying Oil system, where energy
consumption is concentrated in heating and cooling, while dynamic electrical
equipment contributes a small portion.
The reasonable and consistent distribution of energy costs with the capital cost
structure shows that the project has a consistency between equipment design and
operational requirements, while also offering potential for future optimization of
thermal efficiency and fuel cost reduction through heat exchange improvement and
energy reuse strategies.

4. Annual Operating Cost - AOC

AOC=¿ costs+ Variable costs
Where:
● Fixed Costs are items that do not change according to production volume,
including labor, maintenance, asset depreciation, insurance, and administrative
management.
● Variable Costs change according to output, including raw material costs,
consumable chemicals, operational energy, waste treatment, and production
supplies.
Labor Costs
Labor costs can be estimated by:
N OL =¿
Where:
● N OL: Operating Labor per Shift
● P: Particulate steps
● N mp : Non-particulate equipment (Compressors, Towers, Reactors…)
N mp=∑ equipment
N OL =¿

N mp is 16

15
Total operator=¿ N OL × 4.5=4 × 4.5=18
Labor cost=18 ×75 000=1,350,000 USD ∕ year

Raw Material costs can be estimated by:

Raw Material Costs=∑ ( F i ×C i × t)
i

Where:
F i: Flow rate of raw material i (kg/h or kmol/h)
C i: Purchase price of raw material i (USD/kg or USD/kmol)
t : Annual operating time (hour or year)

Raw Material Costs=∑ ( F i × Ci ×t )

i

¿ ∑ ( 1 620.625 ×37 ×8766 )

i

¿ 525,636,753.8 USD ∕ year

AOC=1,350,000 USD / year +525,636,753.8 USD/ year

¿526,986,753.8 USD ∕ year

5. Overall Economic Feasibility

Financial feasibility is evaluated through criteria such as Net Present Value
(NPV), Internal Rate of Return (IRR) and Payback Period (Payback Period). In
particular, NPV is the most popular measure to evaluate the net profitability of a
project throughout the investment life cycle:
N
CF t
NPV =∑ −I 0
t=1 (r +1)t
with CF t is the net cash flow at year t , r is the discount rate (assumed to be 10%), I0 is
the initial investment cost, and N is the number of years of operation (10 years from
2025). IRR is the value of r that makes NPV =0 . The project is considered feasible
when NPV > 0 and IRR>r (required interest rate). Payback period is the number of
years it takes for the accumulated cash flows to cover the initial investment.

16
Revenue = 1,618.851 ×40 × 8766=567,633,914.6 USD ∕ year
CF t = Revenue – AOC =567,633,984.768−526,986,753.8
¿ 40,647,160.84 USD ∕ year
N
CF t
Net Present Value ( NPV ) = ∑( t
−I 0
t =1 r +1 )
10
40,647,160.84
¿∑ −47,893,176.737
t=1 ( 0 ,1+1 )1
¿ 201,866,030.7 USD ∕ year

Additionally, the Profitability Index (PI) can be calculated:

N
CF
∑ (r +1)t t
t =1
PI =
I0
N 10
CF t
∑ (r +1)t ∑ 40,647,160.84
( )1
t =1 t=1 0 ,1+1
PI = = =5.21
I0 47,893,176.737
A project is profitable when PI > 1.

Payback Period:
I 0 47,893,176.737
Payback period= = =1.18 year
CF t 40,647,230.97

6. Net Present Value (NPV)

The project has high economic feasibility with NPV = +153,973 million USD,
return on investment exceeding the cost of capital (PI = 5.21) and payback period of
only 1.18 years. Large annual net cash flow (40,647 million USD) and reasonable cost
structure (focus on main equipment, energy optimization) are key factors. Risks
mainly come from fluctuations in raw material prices and depend on continuous
operation.

17
VII. CONCLUSION AND RECOMMENDATION
The project is considered economically feasible based on three standard financial
indicators - Net Present Value (NPV), Internal Rate of Return (IRR), and Payback
Period- all of which demonstrate strong potential for development. However, issues
related to health, safety, and the environment (SHE) have not yet been thoroughly and
clearly assessed. A notable concern involves the distillation system, specifically the
design that aims to separate three products simultaneously, as well as the recovery of
acetic acid vapor through condensers or filtration units. While the operation of such
equipment is theoretically feasible, in practice, it requires advanced and detailed
engineering designs involving complex technological processes. Therefore, a
comprehensive analysis of the design procedures is essential to ensure optimal and
stable production performance.

18
VIII. REFERENCES
1. James M. Douglas, Conceptual design of chemical processes, McGraw-Hill book,
1988
2. Richard Turton, Richard C. Bailey, Wallace B. Whiting et al., Analysis, synthesis,
and design of chemical processes, Pearson Education Inc, May 2012, pp. 1214 – 1220.
3. Smith, J. M., Chemical Engineering Kinetics, 3rd ed. (New York: John Wiley and
Sons, 1981), 224–228.
4. Towler & Sinnott, Chemical Engineering Design, 2nd Ed., 2013, page 339.

19
IX. APPENDICES
1. Appendix: Process flow diagram

Figure 1: Process flow diagram of Drying Oil Production

2. Appendix: HYSYS diagram

Figure 2: HYSYS diagram of Drying Oil Production

20
 3. Equipment list

Number Reflux Condenser Reboiler Condenser Reboiler Condenser Reboiler

of Trays Ratio Temp. Temp. P (kPa) P (kPa) Duty (kJ/h) Duty (kJ/h)

T-100 10 1 133.5 362.6 122.2 136 1.853E+006 5.343E+006

T-101 10 1 119.7 260.8 108.5 122.2 4.975E+005 9.070E+005

Table 1: Distillation Columns Data

Pressure
Pressure Drop (kPa) Tempurature
(kPa)

F-100 136.0 12 175

Table 2: Separator Data

Pressure (kPa) Temperature Duty (kJ/h) ACO Conversion (%)

R-100 195 334.4 0.000E-01 17.83

Table 3: Reactor Data

Table 4: Heaters, Condensers and Reboilers Data

Type Area (m2) Duty (kJ/h)

E-100 Heater 5.485E+006

E-101 Cooler 4.292E+006

E-102 Condenser 1.853E+006

E-103 Reboiler 5.343E+006

E-104 Condenser 4.975E+005

E-105 Reboiler 9.070E+005

E-106 Cooler 3.743E+006

Power (kW)

P-100 0.9190
21
P-101 3.435E-002
 Table 5: Power of pumps

4. Economics analysis

Unit Cost Energy Cost

Equipment Duty/Power Power (kW)
(USD/kWh) (USD/year)

E-100 5,485,000 kJ/h 1523,61 0,035 $467.459

E-101 4,292,000 kJ/h 1192,22 0,004 $41.804

E-102 1,853,000 kJ/h 514,72 0,004 $18.048

E-103 5,343,000 kJ/h 1484,17 0,035 $455.358

E-104 497,500 kJ/h 138,19 0,004 $4.845

E-105 907,000 kJ/h 251,94 0,035 $77.298

E-106 3,743,000 kJ/h 1039,72 0,004 $36.457

T-100 Condenser 1,853,000 kJ/h 514,72 0,004 $18.048

T-100 Reboiler 5,343,000 kJ/h 1484,17 0,035 $455.358

T-101 Condenser 497,500 kJ/h 138,19 0,004 $4.845

T-101 Reboiler 907,000 kJ/h 251,94 0,035 $77.298

P-100 0.9190 kW 0,919 0,07 $564

P-101 0.03435 kW 0,03435 0,07 $21

Table 6: Energy Cost Table

Type Energy Cost (USD/year)

22
 Fuel/Steam $88.515,649

Cooling Water $1.101.408,102

Electricity $21,078

Table 7: By Energy Type

Base size
Equipment Actual size (S) Unit C₀ (USD) n Scaled cost (USD)
(S₀)

P-100 0,919 5 kW 14587,32 0,6 $5.279,379

E-100 1523,61 100 m² 69698,4 0,6 $357.229,320

E-101 1192,22 100 m² 69698,4 0,6 $308.344,815

X-100 5 2,83 m³ 54278,4 0,6 $76.372,530

T-100
1484,17 0,93 m 101772 0,55 $5.878.664,758
(Reboiler)

E-102 514,72 100 m² 69698,4 0,6 $186.279,659

E-103 1484,17 100 m² 69698,4 0,6 $351.651,924

T-101
251,94 0,93 m 101772 0,55 $2.216.542,722
(Reboiler)

E-104 138,19 100 m² 69698,4 0,6 $84.626,967

E-105 251,94 100 m² 69698,4 0,6 $121.339,188

E-106 1039,72 100 m² 69698,4 0,6 $284.035,516

Mix-101 3,5 2,83 m³ 46260 0,6 $52.550,227

CRV-100 3,5 2,83 m³ 23130 0,6 $26.275,113

RCY-1 1,2 5 kW 14587,32 0,6 $6.195,869

E-10D 350,5 100 m² 69698,4 0,6 $147.922,971

P-101 0,03435 5 kW 14587,32 0,6 $734,767

Total Purchase Equipment Cost (PEC) $10.104.045,725

23
 Total Installed Cost (TIC) = PEC x Lang factor = PEC x 4.74 $47.893.176,737

Table 8: Equipment Scaling Data

Type Scaled cost (USD)

Pump centrifugal $12.210,015

Shell & Tube HX $1.841.430,361

Vertical Vessel $76.372,530

Distillation Column $8.095.207,479

Other $78.825,340

Table 9: Cost by Equipment Type

Year Initial Investment (I₀) Net Cash Flow (CFₜ) Discounted Cash Flow (10%)

0 -47.893 -47.893 -47.893

1 0.000 40.647 36.952

2 0.000 40.647 33.593

3 0.000 40.647 30.539

4 0.000 40.647 27.762

5 0.000 40.647 25.238

6 0.000 40.647 22.943

7 0.000 40.647 20.857

8 0.000 40.647 18.961

9 0.000 40.647 17.237

10 0.000 40.647 15.669

SUM -47.893 406.470 201.866

Table 10: The project pays off quickly after just over 1 year of operation

24
Figure 3,4: Preliminary design and information of the flows through the distillation
column T-100

25
Figure 5,6: Preliminary design and information of the flows through the distillation
column T-101

26