LUND UNIVERSITY DEPT. OF CHEMICAL ENGINEERING KET050 FEASIBILITY STUDIES ON INDUSTRIAL PLANTS FINAL REPORT ON PURIFICATION OF GREEN METHANOL FROM PULP PRODUCTION PRESENTED TO INVICO METHANOL PRINCIPAL INVESTIGATORS: RISHI MIDHA, KALEY OGILVY, TAYLOR SKINGLE, CAROLYN BUCHANAN, GRAHAM MANTAY TUTOR: OLA WALLBERG JUNE/2015
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LUND UNIVERSITY
DEPT. OF CHEMICAL ENGINEERING
KET050 FEASIBILITY STUDIES ON INDUSTRIAL PLANTS
FINAL REPORT
ON
PURIFICATION OF GREEN METHANOL
FROM PULP PRODUCTION
PRESENTED TO
INVICO METHANOL
PRINCIPAL INVESTIGATORS:
RISHI MIDHA, KALEY OGILVY, TAYLOR SKINGLE, CAROLYN BUCHANAN, GRAHAM MANTAY
TUTOR:
OLA WALLBERG
JUNE/2015
i
Abstract
Invico Methanol, an engineering company, works to provide different methods of creating
βgreenβ energy. Invico has recently created a process that utilizes liquid-liquid extraction of raw
methanol in order to remove compounds containing sulfur as well as recover the solvent and
methanol from the extraction. Invico has successfully verified parts of this process in pilot scale.
However, the aim is to design full-scale equipment.
The purpose of this project was to complete an in-depth analysis and perform a technical-
economical comparison of the extraction and purification of methanol from a pulp mill.
Methanol that can be exported from a kraft pulp mill can be used for bio-diesel production, thus
lowering the environmental impact of the plant through creating a biofuel. This report covers
various issues including process designs of our suggested process and other competing
processes. Our recommended design consists of a methanol cleaning system with two
distillation columns, a liquid-liquid extraction column, and an oil cleaning system for cost
reduction, which is achieved using a stream stripper. Each of these designs were simulated and
evaluated with Aspen Plus V8.2 software, which solved the material and energy balances
involved in the process. An investment and operating cost evaluation was also completed, and
the total determined fixed plant cost is 61.3 million SEK and the total operating annual cost is
43.0 million SEK. Lastly, a sensitivity analysis was completed to understand how the selling price
of the reduced methanol, the cost of extraction oil, the interest rate, and maintenance costs
impacted the plant value. The sensitivity analysis showed that at the expected interest rate and
maintenance costs of 11% and 6% (of total plant cost) respectively, the value of the plant would
not be damaged too greatly. More importantly, it was determined that both the selling price of
methanol and the cost of the extraction oil have the power to make this project completely
viable or quite unviable. It is expected that since biofuels are not yet in very high demand, and
the quality of the produced biofuel is not likely to meet the standards of regular methanol, the
selling price of the methanol is likely to be below average and have a significant negative impact
on profits. However, due to the oil cleaning system, if a low-cost extraction oil is found and used
for this process, it has the capability to offset the expected losses from the low selling-point of
the methanol, especially for the first few βtestingβ years of the project.
Final recommendations would include much more detailed research with regards to variable
costs and selling prices, and the investigation of a water cleaning system for the steam and water
waste streams, as if they are reused, variable costs could be reduced significantly.
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Table of Contents
ABSTRACT ........................................................................................................................................ I
TABLE OF CONTENTS ....................................................................................................................... II
LIST OF TABLES .............................................................................................................................. III
The Pulping Process.................................................................................................................................................................. 1
The Kraft Process ....................................................................................................................................................................... 2
PROCESS OVERVIEW ....................................................................................................................... 3
Oil Cleaning System ................................................................................................................................................................... 9
Mixer and Pump ...................................................................................................................................................................... 17
SENSITIVITY ANALYSIS AND FEASIBILITY STUDY ............................................................................. 24
Methanol Sale Price ................................................................................................................................................................ 26
Data Tables ................................................................................................................................................................................ 41 Plant Costs (Ulrich Estimations) ........................................................................................................................................................ 41 Stream Conditions .................................................................................................................................................................................... 44 Unit Operating Conditions .................................................................................................................................................................... 49 Economic Tables and Investment Calculations........................................................................................................................... 51
List of Tables
Table 1: Feed conditions of the raw material to purified, and the physical properties of each component _________ 1 Table 2: Stream conditions before and after the first distillation column, with stream S1 representing the incoming vapour stream, S2 representing the liquid top stream, and S3 representing the liquid bottoms stream ____________________________________________________________________________________________________________________ 5 Table 3: Stream conditions before and after the extraction column, with stream S5 representing the incoming diluted liquid stream, S8 representing the incoming oil (n-hexadecane) stream, S9 representing the oil-heavy top stream, and S10 representing the refined bottoms stream. _______________________________________________________ 6 Table 4: Stream conditions before and after the second distillation column, with stream S10 representing the incoming purified liquid stream S11 representing the final product stream to be shipped to storage, and S12 representing the a waste stream of primarily water. __________________________________________________________________ 7 Table 5: Detailed Aspen results of the final product stream ___________________________________________________________ 8 Table 6: Stream conditions before and after the steam striper, with stream S20 representing the incoming contaminated oil-heavy liquid, STEAMSTR representing the pure steam stream used for stripping, STRIPTOP representing the waste vapour stream, and STRIPBOT representing the cleaned oil, to be recycled back to the extraction column ______________________________________________________________________________________________________ 10 Table 7: NRTL model binary parameters for Ξ±-pinene (1) and water (2). The first set of data is extracted from a paper on liquid-liquid equilibrium for binary mixtures (Utami, Sutijan, Roto, & Sediawan, Liquid-Liquid Equilibrium for the Binary Mixtures of Alpha-Pinene+Water+Alpha-Terpineol+Water, 2013). The second set of data is extracted from a paper on liquid-liquid equilibrium for three component systems (Utami, Sutijan,
iv
Roto, & Sediawan, Liquid-Liquid Equilibrium for System Composed of Alpha-Pinene, Alpha-Terpineol and Water, 2013). ___________________________________________________________________________________________________________ 11 Table 8: Summary of Distillation Column Design Parameters _______________________________________________________ 14 Table 9: List of variables used in Equation 4, their meaning, and units _____________________________________________ 15 Table 10: Summary of Steam Stripper Design Parameters ___________________________________________________________ 16 Table 11: Determined power requirements for the two inline mixer and the centrifugal pump, based off of Aspen results and the flow being handled (Silverson, 2015) _________________________________________________________ 17 Table 12: Summary of Heat Exchanger Design Parameters (Teralba, 2015) _______________________________________ 18 Table 13: Ulrich Cost Estimation for Distillation Column 1 (EconExpert, 2000) ____________________________________ 19 Table 14: Summary of fixed plant installation and equipment cots, and a final plant cost _________________________ 19 Table 15: Average Heating Oil Costs ___________________________________________________________________________________ 20 Table 16: Summary of operating costs of plant employees ___________________________________________________________ 21 Table 17: Estimation of income from methanol production __________________________________________________________ 21 Table 18: Summary of storage costs for products and feed __________________________________________________________ 23 Table 19: Total annual operating costs for the plant _________________________________________________________________ 24 Table 20: Factors Affecting Net Present Value ________________________________________________________________________ 25 Table 21: Summary of Reference NPV Conditions ____________________________________________________________________ 26 Table 22: Summary of Costs____________________________________________________________________________________________ 34 Table 23: Physical parameters of the two phases in the extraction column (Sigma Aldrich, 2014) ________________ 40 Table 24: Ulrich cost estimation for Distillation Column 2 (EconExpert, 2000) _____________________________________ 41 Table 25: Ulrich cost estimation for Inline Mixer 1 and 2 (EconExpert, 2000) ______________________________________ 42 Table 26: Ulrich cost estimation for the centrifugal pump (EconExpert, 2000) _____________________________________ 42 Table 27: Ulrich cost estimation for Heat Exchanger B13 (EconExpert, 2000) _____________________________________ 43 Table 28: Ulrich cost estimation for Heat Exchanger B6 (EconExpert, 2000) _______________________________________ 43 Table 29: Ulrich cost estimation for the Steam Stripper (EconExpert, 2000) _______________________________________ 44 Table 30: Ulrich cost estimation for the Extraction Column (EconExpert, 2000) ___________________________________ 44 Table 31: Detailed stream conditions for S1-S12 _____________________________________________________________________ 45 Table 32: Detailed stream conditions for S14-WATER2 ______________________________________________________________ 47 Table 33: Summary of Factors affecting NPV _________________________________________________________________________ 51 Table 34: NPV and Pay off time calculation summary ________________________________________________________________ 52
1
Introduction The purpose of this project is to explore the possibility of cleaning the effluent from the Kraft
pulping process, by isolating methanol and ethanol from water, Ξ±-pinene, acetone and various
sulfur-containing compounds. The recovered methanol and ethanol can be used to produce bio-
diesel, an alternative fuel source, potentially used for transportation or as fuel for the kraft
pulping process. This area of study is valuable as the demand for eco-friendly alternatives to
fossil fuel is growing rapidly. For this project to be successful the resulting methanol must meet
industrial standards with regards to sulphur content, and the operating/investment costs of the
process must be justifiable. The pulp mill effluent content has been pre-determined, and is given
by Table 1, which also includes physical properties of the effluent components.
Table 1: Feed conditions of the raw material to purified, and the physical properties of each component
Chemical Methanol Ethanol Turpentine Acetone DMS DMDS Ammonia Water
Percent of Feed, by Weight
(%)
78.8 3.1 12.1 0.8 0.4 0.3 0 4.6
Boiling Point (Β°C)
64.7 78.37 90-115 56 37 110 -33.34 100
Molecular Weight (g/mol)
32.04 46.07 136 58.08 62.1 94.20 17.03 18.02
Density (kg/m3)
791.80 789.00 854-868 791.00 846 1060 0.73 999.97
The industrial standard, generally speaking, requires the water content of the methanol product
to be 0.1%, by mass, or less, and the sulfur content to be reduced to 10 ppm or less. The
turpentine must be completely removed.
In this report, existing methanol cleaning techniques will be evaluated, compared, and utilized to
formulate a potentially unique process for methanol purification. To analyze the viability of the
formulated design, an Aspen model will be created, and dimensioning, cost analysis, and
feasibility studies will be conducted.
The Pulping Process
The pulping process is a method of turning materials from wood products (sawmill residue, logs
and chips, recycled paper) into pulp, which can be delivered to a paper mill for further
processing. Wood-based materials entering a pulp mill consist of cellulose fibers, lignin, and
hemicelluloses. The pulp product primarily consists of cellulose fibers, which are the desirable
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component for paper production. Lignin serves the purpose of binding the cellulose fibers
together. In its simplest form, the pulping process is the act of degrading the lignin and the bulk
structure of the incoming wood material, leaving behind pure cellulose fiber material. This can
be achieved by mechanical or chemical means. The mechanical method involves physically
ripping wood fibers apart. The general chemical method involves chemically breaking down the
structure of the lignin and hemicellulose into a water-soluble compound, which can then be
washed away from the cellulose fibers. Over 80% of global pulp production is achieved using the
kraft method, which is a chemical process (European Paper & Packaging Industries, 2015).
The Kraft Process
The chemical degradation of lignin in the kraft process is facilitated by a combination of sodium
sulfide (Na2S) and sodium hydroxide (NaOH). This combination is commonly referred to as
βwhite liquorβ and is the primary source of sulphur-containing compounds throughout the rest of
the process. Figure 1 displays an overview flowchart of the kraft process.
Figure 1: General scheme of the kraft pulping process (Steltenkamp, 1985)
The yellow path in Figure 1 is the main process, where wood chips and wood residues are
reduced to market-quality pulp. The flows beneath the yellow process are primarily waste
treatment and recovery operations to boost the efficiency of the pulping process. For example,
white liquor is continually regenerated in a causticizing process, so it can repeatedly be used to
break the lignin-cellulose bond.
The βevaporationβ section in Figure 1 is the most relevant aspect to the content of this report. The
evaporator unit operation takes black liquor heavily diluted by water, and removes a large
fraction of the water, so the black liquor can be combusted to form βsmelt,β or green liquor. The
water that is drawn off by the evaporator contains approximately 1% methanol, and passes
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through a sequence of distillation columns and filters to bring it to the composition given by
Table 1. A box diagram outlining the distillation process is shown by Figure 2.
Figure 2: Post-evaporation recovery of methanol in the kraft process. Addition of H2SO4 during pre-treatment lowers the pH of the stream, removing H2S and CH4S as gasses, and forming an (NH4)2SO4 precipitate. The
stripper column usually contains 20 trays, and the methanol distillation column contains 15 trays.
It can be seen that following the process given by Figure 2, the outlet conditions are the same as
those shown in Table 1. From this point, the stream enters the section of the process that is
designed and evaluated in this report.
Process Overview The proposed process consists of six unit operations, each which plays a role in producing a
purified methanol stream by removing any undesirable components. The initial stream, as
pinene), ammonia, and water. Roughly 0.25 kilograms per minute of effluent are fed into the
system, which translates to around 19 liters per minute. The finalized recommended process to
achieve the goals as outlined in the Introduction section is shown by Figure 3. Throughout this
section, numerous tables are used to display the stream and operating conditions data for the
specified process. All of the detailed raw data tables are shown in the Unit Operations Section
and Stream Conditions Section in the Appendix.
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Figure 3: The suggested purification process
Boiler Before the first unit operation in the formulated process, there is a boiler (not shown in Figure
3). The boiler is not within the design scope of the project; however its purpose is to boil off all
of the components other than the H2SO4 acid in the inlet stream. The resulting acid-free vapor
stream (stream S1) is collected and continues along the process, while the bottom liquid stream
containing mostly acid is discarded from the process.
Distillation Column 1 The next step in the process is the first distillation column with reflux, shown by component B1
in Figure 3. This column has 20 trays, has a bottom stage temperature of 67Β° C, and the top stage
temperature is 46Β° C. The incoming vapour stream (S1) is split into a liquid top and liquid
bottoms stream. The goal of this step is to remove most of the DMS and acetone, which becomes
concentrated in the top vapour stream. The liquid bottoms are sent to a boiler reflux to further
remove DMS and Acetone. The stream circulates in this way until a desirable fraction of DMS and
acetone has been removed. It should be noted that not all acetone is removed in this step,
although a very small portion remains in stream S3. Also, the DMS is practically completely
eradicated in stream S3. Lastly, it should be noted that a small portion of methanol is lost in this
distillation column to the tops liquid stream. This is why there is a reflux condenser at the top
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stream, to minimize the loss of desirable compounds. The boilup ratio and condenser reflux ratio
have been optimized to minimize methanol losses, and to maximize the removal of acetone and
DMS. The technical specifications of Distillation Column 1 are shown in the Unit Operating
Conditions heading in the Appendix section. Table 2 outlines the relevant stream conditions
with respect to Distillation Column 1. The bottoms stream is then passed onto the next step of
the process.
Table 2: Stream conditions before and after the first distillation column, with stream S1 representing the incoming vapour stream, S2 representing the liquid top stream, and S3 representing the liquid bottoms
Ξ±-pinene KMOL/HR 0.8882 trace 0.8878 water KMOL/HR 2.553 trace 2.553
Inline Mixer 1 The bottoms liquid stream from Distillation Column 1 (Stream S3 in Figure 3) is passed to an
inline mixer (component B2 in Figure 3) where it is diluted with water, reducing the stream
from approximately 85% methanol content to 71% methanol content. An inline mixer was
selected to provide sufficient mixing of the two liquid streams, thus forming stream S5, prior to
them entering the extraction column. The addition of water in this step is important as it creates
a significant solubility difference in the stream for the following step, which involves a liquid-
liquid extraction to reduce the concentrations of acetone, DMS, DMDS, and turpentine in the
stream. The technical specifications of Inline Mixer 1 are shown by the Unit Operating
Conditions heading in the Appendix section.
Extraction Column Following the Inline Mixer, Stream S5 is passed into an extraction column, which is also fed with
a stream of lightly diluted n-hexadecane (Stream S8). Stream S8 is the resulting stream from the
oil recycle loop, which will be discussed later in the report. This extraction column has a top
stage temperature of 23.7Β° C and a bottom stage temperature of 19.6Β° C. The goal of this liquid-
liquid extraction is primarily to reduce the concentrations of Ξ±-pinene and DMDS, although it
also further reduces the concentration of acetone and DMS. Table 3 outlines the relevant stream
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conditions with respect to the extraction column. Detailed technical specifications for the
extraction column are shown in the Unit Operations Heading in the Appendix section.
Table 3: Stream conditions before and after the extraction column, with stream S5 representing the incoming diluted liquid stream, S8 representing the incoming oil (n-hexadecane) stream, S9 representing the oil-heavy
top stream, and S10 representing the refined bottoms stream.
It can be seen in Table 3 that a significant reduction of Ξ±-pinene is achieved with the extraction
column. A fairly significant amount of DMDS and acetone are also removed. Trace amounts of
DMS are removed, although the DMS is already very low in concentration. From the resulting
βcleanedβ stream (S10) it can be seen that the stream is close to approaching the desired
purification of methanol and ethanol, however there is still a concentration of DMDS that is
above 10 ppm and the water content is above 0.1%. It is therefore still not adequate. The top
stream from the extraction column (S9) primarily consists of n-hexadecanes, and is sent to an
oil-cleaning system so the oil can be reused within the extraction column. This system is
described in future sections.
The importance of the addition of water before the extraction column can be seen from Table 3,
where the Ξ±-pinene is almost completely eradicated from the stream. The highly polar nature of
water, paired with the highly non-polar nature of n-hexadecane causes the Ξ±-pinene (a non-
polar solvent) to be drawn out of the stream with the exiting liquid stream (S9). In the oil-
cleaning system, Stream S9 will primarily need to be stripped of Ξ±-pinene and DMDS. Another
observation from Table 3 is that trace amounts of ethanol and methanol are lost to stream S9 in
the extraction process. It is important to note, though, that following the oil cleaning, this stream
is returned to the extraction process, therefore some of this ethanol and methanol could be
returned to the product stream. The remaining methanol and ethanol removed from the oil
could be recovered using further processing steps; however this is out of the scope of the work
for this report.
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Distillation Column 2
Stream S10, the incoming stream for the second distillation column, is still relatively
concentrated with water and DMDS, therefore the primary goal of this column (B5) is to reduce
both of these compounds. The distillation column, which has 15 trays, has a top stage
temperature of 62.0Β°C and a bottom stage temperature of 78.9Β°C. The bottom stage temperature
exceeds the boiling point of both methanol and ethanol, thus the concentrated product is
removed in stream S11, the top stream of the distillation column. The bottoms stream, S12,
primarily consists of water, which can be treated and reused, or sent to waste. Table 4 outlines
the relevant stream conditions concerning the second distillation column.
Table 5 contains the detailed results of the final stream of this product. Detailed technical
specifications for Distillation Column 2 are shown in Unit Operations Heading in the Appendix
section.
Table 4: Stream conditions before and after the second distillation column, with stream S10 representing the incoming purified liquid stream S11 representing the final product stream to be shipped to storage, and S12
representing the a waste stream of primarily water.
S10 S11 S12 From B4 B5 B5
Phase: Liquid Liquid Liquid Component Mole Flow Units
Molar Enthalpy CAL/MOL -56460 Mass Enthalpy CAL/GM -1753 Enthalpy Flow CAL/SEC -349200 Molar Entropy CAL/MOL-K -54.83 Mass Entropy CAL/GM-K -1.702 Molar Density MOL/CC 0.02325 Mass Density GM/CC 0.7489
Average Molecular Weight 32.22
The purified product stream, shown by Table 5, has a very concentrated combined methanol and
ethanol mass fraction of 99.2%. According to the standards introduced in the previous section,
the product stream was required to include less than 0.1% water by mass and less than 10 parts
per million (ppm) of sulfur.
Table 5 confirms that the final mass fraction of sulfur-containing compounds is 3.32x10-7, which
is significantly less than 10 parts per million. The sulfur removal requirement is therefore
achieved. The final water concentration is observed to be 0.7%, by mass, slightly higher than the
desired concentration of 0.1%. After various optimization methods with the boilup and reflux
ratios, the team was unsuccessful in achieving the desired water concentration. A decision had
to be made between adding another distillation column and looking for further improvements to
9
the existing column to achieve the desired concentration. Due to the astronomical cost of adding
another distillation column to achieve such a small purification of the final stream, the team
decided that another column would not be added. Instead, it is recommended that further
research and modelling of this column are completed. It is suspected that finely tuning the
column by changing the type of trays, the number of trays and other variables could achieve the
desired separation. This option seems much more cost effective and reasonable compared to
adding in a second column. Unfortunately, the team did not have the expertise or resources
(time and knowledge) to complete this tuning of the column.
Oil Cleaning System
Stream S9, which exits the extraction column (B4), is highly concentrated in n-hexadecane, thus
providing the opportunity for reuse in the extraction column if it is adequately cleaned. The
removal of undesirable compounds (methanol, ethanol, acetone, DMS, DMDS, Ξ±-pinene, water) is
achieved using a steam stripper (B3), which uses steam to remove all of the volatile organic
compounds in the stream and release them as a vapour (Stream STRIPTOP). The resulting
bottoms liquid stream STRIPBOT is almost entirely consisted of n-hexadecane and a small
amount of water. Prior to entering the steam stripper and mixing with high-pressure steam, the
stream passes through a heat exchanger (B6) and a pump (B9) to achieve the desired
temperature and pressure for efficient removal. The technical specifications for the heat
exchanger, pump, and the steam stripper are shown by the Unit Operations Heading in the
Appendix. The STRIPTOP stream can be sent off as a waste to a handling system, and STRIPBOT
can be recycled and used again in the extraction column, after further treatment. As an
alternative, the STRIPTOP stream could be sent to further processing units to recover some of
the methanol and ethanol lost in this stream. This is out of the scope of the project and is simply
a recommendation for further improvements. Table 6 contains all the stream information
regarding the steam stripper.
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Table 6: Stream conditions before and after the steam striper, with stream S20 representing the incoming contaminated oil-heavy liquid, STEAMSTR representing the pure steam stream used for stripping, STRIPTOP representing the waste vapour stream, and STRIPBOT representing the cleaned oil, to be recycled back to the
Table 6 shows that the STRIPBOT stream is primarily n-hexadecane, however the stream
conditions are not quite equivalent to the stream conditions of S8, the oil feed stream for the
extraction column. Thus, stream S16 is implemented to feed a small amount of n-hexadecane to
the recycle stream to account for the minor losses of n-hexadecane during the extraction and
steam stripping. This stream is added to stream STRIPBOT using a second inline mixer, Inline
Mixer 2. Another heat exchanger (B13) is installed before the extraction column to ensure the
temperature of the oil is adequate for liquid-liquid extraction. The operating conditions for each
unit in the oil separation cycle can be found in the Appendix. With this cycle completed, the
entire process is converged and operating within the expected standards.
Decanter Modelling One of the main goals of the process modelling phase of this project was to determine the
viability and functionality of adding a decanter into the process. The purpose of the decanter
was to remove the Ξ±-pinene from the methanol stream prior to the extraction unit. It was
proposed that adding water to the stream would result in a change in polarity of the stream, thus
allowing a phase separation and the resulting removal of Ξ±-pinene in one of the phases, while
keeping methanol in the other. Since Aspen was lacking information regarding the binary
interactions between Ξ±-pinene and water, a literature search was completed to look for
appropriate NRTL model binary interaction parameters for these chemicals (NRTL was the
model used in the Aspen modelling). Two different sets of binary data were found and both were
examined in Aspen. Table 7 shows both sets of parameters that were found and examined.
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Table 7: NRTL model binary parameters for Ξ±-pinene (1) and water (2). The first set of data is extracted from a paper on liquid-liquid equilibrium for binary mixtures (Utami, Sutijan, Roto, & Sediawan, Liquid-Liquid Equilibrium for the Binary Mixtures of Alpha-Pinene+Water+Alpha-Terpineol+Water, 2013). The second set of data is extracted from a paper on liquid-liquid equilibrium for three component systems (Utami, Sutijan, Roto, & Sediawan, Liquid-Liquid Equilibrium for System Composed of Alpha-Pinene, Alpha-Terpineol and Water, 2013).
Parameters Data Set 1 Data Set 2
Bij (J/mol) -5.851e-007 5652.49
Bji (J/mol) 5752.89 7305.75
It was found that the difference between using the first set of parameters and the second set was
minimal. Both produced a ternary diagram for water, methanol and Ξ±-pinene that looked like the
one presented in Figure 4.
Figure 4: Ternary plot for a mixture of water, methanol and Ξ±-pinene.
Careful consideration of the ternary diagram led to the conclusion that the mole fractions of Ξ±-
pinene and water had to be increased in order to achieve separation. Water was added to the
stream and a small separation was eventually achieved. Unfortunately this separation split the
Ξ±-pinene in a way that half went into each exiting stream. This result seemed useless as the
extraction column still had to remove a good amount of Ξ±-pinene afterwards. In order to justify
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adding an entire new unit into the process, the team needed better separation results. To
achieve a meaningful separation however, the mole fractions of the water and Ξ±-pinene had to
be increased to almost 0.5 each. This would cause the stream to land in the upper area of the
ternary diagram where one of the exiting streams contains virtually all of the Ξ±-pinene and the
other contains virtually none. After a series of trial and error optimization steps, it was
determined that an incredibly large amount of water and Ξ±-pinene had to be added to dilute the
stream to the desired concentrations (almost 5000 kg/hour of each chemical had to be added).
Even after the separation was achieved, the small amount of residual Ξ±-pinene in the methanol
stream was almost equal to the amount that was initially in the stream before adding the
chemicals. It was quite obvious that this was not a viable solution β the cost of the distillation
towers later in the process would be astronomical to separate the water from the methanol.
Furthermore, even if the Ξ±-pinene was recycled, further units would be needed to purify the Ξ±-
pinene and mix it back into the methanol stream.
The team also examined the separation quality at different temperatures within the decanter. By
reducing the temperature of the stream entering the decanter, better separations were generally
achieved. Unfortunately, changing the temperature within the decanter still did not have a great
enough impact to achieve a satisfactory separation.
With the issues mentioned above regarding the implementation of the decanter, coupled with
the fact that the extraction unit was able to remove virtually all of the Ξ±-pinene from the
methanol stream, the team decided that the decanter was simply not a logical or practical option
for this process. No further research was put towards implementing the decanter and the team
focused instead on optimizing the rest of the process.
Process Alternatives Chemrec, a Swedish company focused on optimizing existing pulp and paper plants, has
proposed an alternative process to the one proposed in this report. In their process, waste black
liquor is fed into a high temperature gasification unit along with oxygen (COWI, n.d.). This unit
removes water and some undesirable components from the waste black liquor, resulting in
outlet streams of superheated steam, waste green liquor, and strong black liquor (COWI, n.d.).
The strong black liquor is then passed through a carbon filter in order to remove tars and other
undesirable free carbon. This filter must be initially heated to 1260Β°C to activate the carbon,
which is very energy intensive (COWI, n.d.).
From the carbon filter, the stream of black liquor passes through a sour shift unit. In this unit,
the product synthesis gas composition is optimized through the catalytic reaction of CO and CO2
13
in the black liquor to form methanol (COWI, n.d.). This process is also fairly energy intensive, as
the catalyst most commonly used is only active in a temperature range of 200Β°C-500Β°C (COWI,
n.d.). Once this is finished the stream is fed into an acid gas removal unit. This unit uses solvents
such as C16 under high pressure to remove any sulfur compounds and CO2 present (COWI, n.d.).
The final step in the process involves distillation columns. Two columns are used in series to
remove the remaining water and waste products from the bottoms streams and collect purified
methanol off the top streams (COWI, n.d.).
The Chemrec process appears to work well, but its major disadvantage is that it is very energy
intensive, requiring operation at either high temperatures or high pressure in most parts of the
process. This process is also designed for a much larger scale operation than is being dealt with
in this report. Chemrec installations can process up to 4000 tons per day of black liquor, while
the proposed process handles only about 10 tons per day (COWI, n.d.). This process clearly lies
outside of the scope of the project.
Materials and Equipment Sizing
With a process flow and order selected, the next step is to dimension the individual process
components. This is necessary to ensure proper functionality of the plant, and to allow for cost
estimations of the plant in future sections. In this section, dimensions for each unit operation are
calculated.
Distillation Column Sizing The main design parameters that affect the cost of the two distillation columns are height, inside
diameter, number of trays, and operating pressure (Vasudevan & Ulrich, 2000). The Aspen
simulations automatically calculated the number of trays and operating pressure, while the
height and diameter were calculated using industrial sizing equations. Equation 1 is used to
estimate the height of a distillation column.
π»π = ππ‘ β π‘ (1)
In Equation 1, Nt refers to the number of trays in the column, π‘ refers to the tray spacing, and Hc
refers to the height of the column. Aspen modeling provided the appropriate number of trays,
and industry rules of thumb were used for tray spacing (Separation Technology, 2012), allowing
for the height of the tank to be calculated. The Sample Calculations section shows an example of
the column height calculation.
To calculate the diameter of the distillation tank, a few more steps are required. The diameter of
the tank is calculated based on the flooding characteristics and rules of thumb. The column
14
diameter must be kept high enough relative to the vapour velocity within the column to prevent
flooding (Price, 2003). Also, the height to diameter ratio of the tank should never exceed 30 and
is ideally kept below 20 (Price, 2003). The tank diameter was first calculated based on flooding
characteristics, and then checked against the rules of thumb to ensure the height to diameter
requirement was met. Calculations were done for the distillation column to operate at a vapour
velocity of 85% of the flooding capacity. Equation 2 shows how the vapour velocity was
calculated (Hoogstraten & Dunn, 1998).
π’π = πΎποΏ½ππβππ£ππ£
(2)
In Equation 2, π’π refers to the vapour velocity, and ππ and ππ£ refer to the liquid and vapour
densities, respectively. Kl is a constant parameter given in ft/s. The Sample Calculation section
in the Appendix contains exact information on how the vapour velocity was calculated. Given the
design vapour velocity, the cross sectional area of the column was calculated using Equation 3.
π΄ = ππ£0.85π’π
(3)
In Equation 3, A refers to the cross sectional area of the column, and Vv refers to the maximum
volumetric flow rate of the vapour (as calculated by Aspen). Once the cross sectional area has
been calculated the radius can be determined. For Sample Calculations, see Appendix.
Table 8 shows a summary of the design parameters for the two distillation columns. All units
have been converted to metric for clarity.
Table 8: Summary of Distillation Column Design Parameters
Parameter Column 1 Column 2 Number of Trays, N 20 15
Design Vapour Velocity, ππ (m/s)
2.47 0.98
Column Height, Hc,
(m) 12 9
Column Diameter, Dc, (m)
0.55 0.76
Height: Diameter Ratio 2.92 11.81
As shown by Table 8, the height to diameter ratio is well below 20 for both of the distillation
columns, which mitigates the risk of flooding. Refer to the Sample Calculations section for a
more detailed explanation of the given results.
15
Extraction Column Sizing
The dimensioning of an extraction column is based on theory concerning mass transfer
coefficients. The type of column chosen to be used and dimensioned has continuous differential
contact, with countercurrent phase flow where one flow is continuous and the other is
dispersed. The industrial column is designed by determining the geometrical dimensions
(diameter and height), which are based on several factors. The column diameter depends on the
processing capacity and the flooding capacity and can be calculated using Equation 4.
In Equation 5, the Cfair constant was found in literature to be 0.15 m/s for a water/oil mixture
(Separation Technology, 2012). For the densities, ΟL represents the density of water, and ΟV
16
represents the oil density (in kg/m3), which is always less than water. Lastly, Ο represents the
surface tension of the phase in question (in dyn/cm). With voil an3d vwater determined, the
diameter of the column can be found using Equation 4.
After substituting all the values in to Equation 4, the column diameter was calculated to be 4 m.
This unit is designed to be vertically oriented with tower packing. The tower packing material
consists of random polymeric and the vessel material is carbon steel. To see exactly how the
numerical results were generated for the extraction column design, refer to the Sample
Calculations section.
Steam Stripper For the purposes of sizing and cost estimation, it was assumed that the design equations for
Distillation Columns could also be applied to steam strippers. This was deemed to be a fair
assumption as steam strippers in industry are sized similarly to distillation columns, where the
height is a function of the number of trays and the diameter is chosen to prevent flooding (KLM
Technology Group, 2011). The diameter of distillation columns is typically chosen so that the
maximum vapour velocity is between 50-80% of the flooding velocity (Redel, Novoa, Goldina, &
Englert, N/A), while the diameter of strippers are chosen so that the vapour velocity is between
60-80% of the flooding velocity (KLM Technology Group, 2011). This assumption was also
deemed to be safe from an economic standpoint, as the Ulrich cost estimation database does not
distinguish between distillation and stripping columns for the purposes of cost estimation. Table
10 summarizes the design parameters calculated for the steam stripper. The Sample
Calculations section shows the specific method of calculation to obtain these results, which
follow an identical method as the distillation column calculations.
Table 10: Summary of Steam Stripper Design Parameters
Parameter Value Number of Trays, N 20
Design Vapour Velocity, ππ (m/s)
0.35
Column Height, Hc,
(m) 26
Column Diameter, Dc, (m)
2.78
Height : Diameter Ratio 9.35
17
Mixer and Pump For the pump, a power requirement value was generated by Aspen based off the process flow
that the pump could be handling. For the inline mixers, literature values were used to match a
power rating with the flows that the pieces of equipment were likely to handle. Table 11 shows
the mixer and pump results.
Table 11: Determined power requirements for the two inline mixer and the centrifugal pump, based off of Aspen results and the flow being handled (Silverson, 2015)
Equipment Power Requirement (kW)
Inline Mixer 1 1.5 Inline Mixer 2 22.37
Centrifugal Pump 1.51
Heat Exchangers
The design parameters of interest for heat exchangers as far as cost estimation and process
planning are the operating pressure, heat duty, and heat transfer surface area. The Aspen
simulations performed the calculations for the operating pressures and heat duties
automatically, but did not calculate the surface area. The required surface area was calculated
using Equation 6.
π = ππ΄Ξπ (6)
In Equation 6, Q refers to the heat duty, U refers to the heat transfer coefficient, A refers to the
surface area of the exchanger, and the ΞT refers to the change in temperature of the product
entering and exiting the heat exchanger. The temperature change and heat duty can be taken
from Aspen simulations, and the heat transfer coefficient can be estimated from tabulated
literature data, assuming the heating fluid to be steam and the product being heated to have heat
transfer characteristics similar to the majority component of the stream. For example, in the first
heat exchanger, where the fluid being heated is mostly n-hexadecane, the heat transfer
coefficient was estimated to be the tabulated coefficient for heat transfer between n-hexadecane
and steam.
Given the heat duty, heat transfer coefficient, and temperature change, the required surface area
was calculated. Table 12 summarizes the design parameters recommended for the heat
exchangers in the process. The simple calculation for surface area is shown in the Sample
Calculation section of the Appendix.
18
Table 12: Summary of Heat Exchanger Design Parameters (Teralba, 2015)
Parameter Exchanger 1 Exchanger 2 Operating Pressure, P
(bar) 1 1
Heat Duty, Q (kW)
1076 1200
Temperature Change, ΞT 108 116
Heat Transfer Coefficient, U (W/m2K)
1000 1000
Surface Area, A (m2)
9.96 10.34
Investment Cost Analysis
Plant Cost
With the operating parameters and dimensions of the plant equipment confirmed, it is possible
to calculate the cost of each unit operation and thus an estimated final cost for the plant
production. The results represent an estimation of the total βfixed costβ involved with this
process. To complete these calculations, the Ulrich method is used, as it is the most
comprehensive with respect to the factors used to arrive at a final cost. The EconExpert program
utilizes the Ulrich method and was the main source of plant cost estimations, based off of inputs
determined in the previous section. For the first distillation column, Table 13 displays the
EconExpert input and the cost results for the first distillation column.
Operating Cost After determining the capital plant costs, it is possible to determine the cost of operation for the
plant on a per year basis. The total operating cost would include a number of parameters such as
feedstock, maintenance, employees, research and development. The first step was to determine
DISTILLATION COLUMN 1 Cost Summary The cost index is 579.7 Process Vessels (including towers) : Vertically oriented : With sieve-trays Total purchased cost = $ 33365 Material factor = 1.00 Trays 20 Material factor for tower packing = 2.20 Height 12 The bare module cost of tower packing is = $ 12475 Diameter 0.55 The bare module cost is = $ 155470 Pressure 1 Total Bare Module Cost = $ 155470 Contingency and Fee = $ 27985 Total Module Cost = $ 183455 Auxiliary Facilities = $ 55037 Grass Roots Capital = $ 238492
20
the cost of the feedstock, which includes water used for steam and heating oil costs. The cost of
steam was found to be 66 SEK/ton, which equates to 158.4 SEK/day and then multiplied by the
number of days in a year to find the total yearly cost. Similarly, to find the yearly cost of heating
oil, the average cost in US dollars was found, and inflated by multiplying it by 1.5 to account for
error in prices, as Swedish costs were not readily available. Then the actual cost of heating oil
was multiplied by the flow rate of 0.0058 π3/πππ and converted to costs per year in Swedish
currency. The results of all of these calculations are summarized in Table 15.
After all the operating costs have been calculated, the cost of research and development as well
as licensing fees per year must be determined. This is simply just a percentage of the total of the
rest of the operating costs. In the case of research and development it is 1.5% of the other costs,
and in terms of licensing fees it is 3% of these costs. All the costs calculated in terms of operation
of plant can be found summarized below in Table 19.
24
Table 19: Total annual operating costs for the plant
Type of Cost Cost (SEK) Feedstock
Water 58,000 Oil 28,000,000
Maintenance 3,800,000 Spare Parts 570,000 Operators 6,400,000 Supervisors 950,000 Lab Work 950,000 Storage Costs 710,000 Research and Development 620,000 Licensing Fees 1,200,000 Total Operating Cost 43,000,000
Thus, the total plant operating cost is 43,010,647 SEK/year or 5,125,643.32 USD/year.
Sensitivity Analysis and Feasibility Study The main factors affecting the Net Present Value (NPV) for this project could be divided into
three categories: those affecting the Grassroots Capital cost, those affecting the operating costs,
and those affecting the income. The initial list of factors affecting the cost can be found in Table
33 of the Appendix.
The main factors affecting Grassroots capital investment had to do with the sizing and design
parameters of the unit operations in the process. Naturally there were some parameters that
could not be adjusted, as doing so would prevent the process from achieving the desired
performance. It was found, however, that the tray spacing could be adjusted without reducing
the performance of the process, and that doing so would affect the size and cost of the
distillation columns.
The other main sources of uncertainty were in the operating costs and income estimates. The
cost of extraction oil had a range of possible values. Also, the sale value of the final methanol
product and the interest rate over the productβs lifetime were uncertain. Many other operating
costs were calculated by industry rules of thumb provided by the project supervisor. These rules
of thumb came with inherent uncertainty, and as a result also affected the NPV. Table 20 shows
25
how each variable affected NPV for a 20-year economic lifetime, in descending order of how
much each variable affected the NPV.
Table 20: Factors Affecting Net Present Value
Factor Min Value Max Value Change in NPV (SEK) Methanol Sale Price (USD/gal) 0.75 2.53 509 233 516.06
EconExpert. (2000). UlrichDesign. Retrieved 2015, from http://www.ulrichvasudesign.com/econ.html
European Paper & Packaging Industries. (2015). Types of Pulping Processes. Retrieved 2015, from Paper Online: http://www.paperonline.org/paper-making/paper-production/pulping/types-of-pulping-processes
Exchange-Rates. (2015). Swedish Kronor (SEK) to US Dollar (USD). Retrieved 2015, from Exchange-Rates: http://www.exchange-rates.org/Rate/SEK/USD/1-3-2014
Hoogstraten, C., & Dunn, K. (1998). Separation Processes. Retrieved 2015, from McMaster University: http://learnche.mcmaster.ca/wiki_4M3/images/c/c7/Distillation-project-report-vanHoogstraten-Dunn-UCT-1998.pdf
ICIS. (2006, August 28). Indicative Chemical Prices A-Z. Retrieved May 20, 2015, from http://www.icis.com/chemicals/channel-info-chemicals-a-z/
ICIS. (2015). Indicative Chemical Prices A-Z. Retrieved 2015, from ICIS: http://www.icis.com/chemicals/channel-info-chemicals-a-z/
KLM Technology Group. (2011). Distillation Column Selection and Sizing. Retrieved 2015, from KLM Technology Group: http://kolmetz.com/pdf/EDG/ENGINEERING%20DESIGN%20GUIDELINES%20-%20distillation%20column%20-%20Rev%2004%20web.pdf
Koncseg, C., & Barbulescu, A. (2011). Liquid-Liquid Extraction With and Without a Chemical Reaction. Retrieved 2015, from InTech: http://cdn.intechopen.com/pdfs-wm/13535.pdf
Price, R. (2003). Distillation. Retrieved 2015, from http://facstaff.cbu.edu/rprice/lectures/distill7.html
Redel, M., Novoa, A., Goldina, T., & Englert, M. (N/A). Distillation. Retrieved 2015, from University of Utah: http://www.che.utah.edu/~ring/Design%20I/Articles/distillation%20design.pdf
Sigma Aldrich. (2014). Benzaldehyde. Retrieved Febrary 24, 2014, from http://www.sigmaaldrich.com/catalog/product/aldrich/w212717?lang=en®ion=CA
Silverson. (2015). High Shear In-Line Mixers - Technical Information. Retrieved 2015, from Silverson: http://www.silverson.com/us/products/in-line-mixers/technical-information
37
StatsSkuld. (2013). The Statistical wage for Wood chip plant operators, forestry in Sweden. Retrieved 2015, from StatsSkuld: http://en-sv.statsskuld.se/salary/for-Wood+chip+plant+operators%2C+forestry-158027/2013/
Steltenkamp, M. S. (1985, March 26). Patent No. US4507172 A.
Teralba. (2015). Mixing and Agitating Solutions. Retrieved 2015, from Teralba: http://www.teralba.com/heat-exchangers/mixers-and-agitators-solutions
Thermopedia. (2011). Heat Transfer Coefficient. Retrieved 2015, from Thermopedia: http://www.thermopedia.com/content/841/
US Energy Information Administration. (2015, April 1). Petroleum and Other Liquids. Retrieved May 20, 2015 , from http://www.eia.gov/petroleum/heatingoilpropane/
Utami, H., Sutijan, Roto, & Sediawan, W. B. (2013). Liquid-Liquid Equilibrium for System Composed of Alpha-Pinene, Alpha-Terpineol and Water. International Journal of Chemical Engineering and Applications, 21-25.
Utami, H., Sutijan, Roto, & Sediawan, W. B. (2013). Liquid-Liquid Equilibrium for the Binary Mixtures of Alpha-Pinene+Water+Alpha-Terpineol+Water. World Academy of Science, Engineering and Technology, 917-920.
Vasudevan, P., & Ulrich, T. (2000). An Expert System for Capital Cost Estimation. Retrieved 2015, from EconExpert: http://www.ulrichvasudesign.com/cgi-bin/cgiwrap.cgi/econ/econnew.pl
Zygula, T. (2007). A Design Review of Steam Stripping Columns for Wastewater Service. Retrieved 2015, from AIchE: http://www.klmtechgroup.com/PDF/Articles/articles/Steam%20Stripping%20Paper%20Version%20Final.pdf
38
Appendix
Sample Calculations
Distillation Column Design Procedure Calculations were done for the distillation column to operate at a vapour velocity of 85% of the
flooding capacity. Equation 2 shows how the vapour velocity was calculated (Hoogstraten &
Dunn, 1998).
π’π = πΎποΏ½ππβππ£ππ£
(2)
In Equation 2, π’π refers to the vapour velocity, and ππ and ππ£ refer to the liquid and vapour
densities, respectively. Kl is a constant parameter given in ft/s. Kl was calculated using
correlations from literature (Hoogstraten & Dunn, 1998), shown in Equation 3.
πΎπ = 0,26π‘β0,029π‘2
οΏ½1+6πΉππ£2 π‘0,7498
(3)
In Equation 3, t refers to the tray spacing in feet, and Flv is a dimensionless number also
calculated by correlations (Hoogstraten & Dunn, 1998). The equation below shows how Flv was
Process Vessels (including towers) : Vertically oriented : With sieve-trays Total purchased cost = $ 28152
Material factor = 1.00
Trays 15 Material factor for tower packing = 2.20
Height 9
The bare module cost of tower packing is = $ 19875 Diameter 0.76 The bare module cost is = $ 140525
Pressure 1
Total Bare Module Cost = $ 140525
Contingency and Fee = $ 25294
Total Module Cost = $ 165819
Auxiliary Facilities = $ 49746
Grass Roots Capital = $ 215565
42
Table 25: Ulrich cost estimation for Inline Mixer 1 and 2 (EconExpert, 2000)
INLINE MIXER 1 INLINE MIXER 2 Cost Summary
Cost Summary
The cost index is 579.7 The cost index is 579.7
Mixers : Agitators and Inline Mixers : Inline Mixers : Agitators and Inline Mixers : Inline Total purchased cost = $ 9402 Total purchased cost = $ 26240 The bare module cost is = $ 32907 The bare module cost is = $ 91840
Total Bare Module Cost = $ 32907 Total Bare Module Cost = $ 91840 Contingency and Fee = $ 5923 Contingency and Fee = $ 16531 Total Module Cost = $ 38831 Total Module Cost = $ 108371 Auxiliary Facilities = $ 11649 Auxiliary Facilities = $ 32511 Grass Roots Capital = $ 50480 Grass Roots Capital = $ 140883
POWER 1.5KW
POWER 22.37 KW
Table 26: Ulrich cost estimation for the centrifugal pump (EconExpert, 2000)
Pump 1 Cost Summary
The cost index is 579.7
Pumps : Centrifugal
Total purchased cost = $ 5712
Material factor = 1.90
Pressure factor = 1.00
The bare module cost (incl. electric motor drive) is = $ 27567
Table 32: Detailed stream conditions for S14-WATER2
S14 S16
S17 S20 STEAMSTR
STRIPBOT
STRIPTOP
WATER
WATER2
From B6 B12 B9 B3 B3 To B9 B12 B13 B3 B3 B12 B2 Phase: Liqu
id Liquid
Liquid
Liquid
Vapor Liquid Vapor Liquid
Liquid
Component Mole Flow
METHA-01 KMOL/HR
2.713 0 2.03E-29
2.713 0 2.03E-29 2.713 0 0
ETHAN-01 KMOL/HR
0.08714
0 3.16E-50
0.08714
0 3.16E-50 0.08714 0 0
ACETO-01 KMOL/HR
0.002454
0 8.56E-41
0.002454
0 8.56E-41 0.002454 0 0
DIMET-01 KMOL/HR
5.31E-07
0 3.55E-39
5.31E-07
0 3.55E-39 5.31E-07 0 0
DIMET-02 KMOL/HR
0.02794
0 1.78E-17
0.02794
0 1.78E-17 0.02794 0 0
ALPHA-01 KMOL/HR
0.8875
0 9.52E-06
0.8875
0 9.52E-06 0.8875 0 0
WATER KMOL/HR
0.05116
0 0.5581
0.05116
180 0.5581 179.4 5.550 16.65
N-HEX-01 KMOL/HR
66.19 1.750
66.24 66.19 0 64.49 1.702 0 0
Component Mole Fraction
METHA-01 0.03879
0 3.04E-31
0.03879
0 3.12E-31 0.014676 0 0
ETHAN-01 0.001245
0 4.73E-52
0.001245
0 4.85E-52 0.000471 0 0
ACETO-01 3.51E-05
0 1.28E-42
3.51E-05
0 1.32E-42 1.33E-05 0 0
DIMET-01 7.59E-09
0 5.32E-41
7.59E-09
0 5.46E-41 2.87E-09 0 0
DIMET-02 0.000399
0 2.66E-19
0.000399
0 2.73E-19 0.000151 0 0
ALPHA-01 0.012685
0 1.43E-07
0.012685
0 1.46E-07 0.004799 0 0
WATER 0.000731
0 0.008356
0.000731
1 0.008580 0.970679 1 1
N-HEX-01 0.946112
1 0.991643
0.946112
0 0.991418 0.009208 0 0
Component Mass Flow
METHA-01 KG/HR 86.95918
0 6.51E-28
86.95918
0 6.51E-28 86.95918 0 0
ETHAN-01 KG/HR 4.014457
0 1.45E-48
4.014457
0 1.45E-48 4.014457 0 0
ACETO-01 KG/HR 0.142538
0 4.97E-39
0.142538
0 4.97E-39 0.142538 0 0
DIMET-01 KG/HR 3.30E-05
0 2.21E-37
3.30E-05
0 2.21E-37 3.30E-05 0 0
DIMET-02 KG/HR 2.632198
0 1.68E-15
2.632198
0 1.68E-15 2.632198 0 0
ALPHA-01 KG/HR 120.917
0 0.001297
120.917
0 0.001297 120.9157 0 0
48
WATER KG/HR 0.921692
0 10.05579
0.921692
3242.75 10.05579 3233.616 100 300
N-HEX-01 KG/HR 14989.27
396.309
14999.97
14989.27
0 14603.66 385.6022 0 0
Component Mass Fraction
METHA-01 0.00571
0 4.34E-32
0.00571
0 4.45E-32 0.022681 0 0
ETHAN-01 0.00026
0 9.69E-53
0.00026
0 9.95E-53 0.001047 0 0
ACETO-01 9.3E-06
0 3.31E-43
9.3E-06
0 3.40E-43 3.72E-05 0 0
DIMET-01 2.1E-09
0 1.47E-41
2.1E-09
0 1.51E-41 8.61E-09 0 0
DIMET-02 0.0001
0 1.12E-19
0.00017
0 1.15E-19 0.000686 0 0
ALPHA-01 0.00795
0 8.64E-08
0.00795
0 8.88E-08 0.031538 0 0
WATER 6.0E-05
0 0.00066
6.0E-05
1 0.000688 0.843431 1 1
N-HEX-01 0.98582
1 0.99933
0.98582
0 0.999311 0.100577 0 0
Mole Flow KMOL/HR
69.9637
1.75012
66.7990
69.9637
180 65.04891 184.9148 5.5508 16.6525
Mass Flow KG/HR 15204.8
396.309
15010.0
15204.8
3242.75 14613.72 3833.883 100 300
Volume Flow L/MIN 366.978
8.55404
357.800
367.018
78566.6 349.3187 111469 1.6687 5.00617
Temperature C 140 20 128.356
140.102
104.8316 131.0616 140.116 20 20
Pressure BAR 1 1 1 2 1.2 1.044654 0.95 1 1
Vapor Fraction 0 0 0 0 1 0 1 0 0
Liquid Fraction 1 1 1 1 0 1 0 1 1
Solid Fraction 0 0 0 0 0 0 0 0 0
Molar Enthalpy CAL/MOL
-9048
4
-1.0E
06
-9505
0
-9047
0
-57112.8 -94666.39
-56523.36
-68350 -68350
Mass Enthalpy CAL/GM
-416.3
5
-482.
74
-423.0
0
-416.2
9
-3170.242 -421.3811
-2726.219
-3794.0
-3794.03
Enthalpy Flow CAL/SEC
-1.7E0
6
-5314
3
-1.7E0
6
-1.7E0
6
-2855600 -1710500 -2903300 -1.0E05
-316170
Molar Entropy CAL/MOL-K
-345.1
6
-406.
39
-363.0
6
-345.1
2
-9.032567 -362.0526
-11.40727
-39.26 -39.26
Mass Entropy CAL/GM-K
-1.588
2
-1.79
46
-1.615
7
-1.588
0
-0.5013836
-1.611576
-0.55019 -2.1794
-2.1794
Molar Density MOL/CC
0.0031
0.0034
0.0031
0.0031
3.82E-05 0.0031036
2.76E-05 0.0554 0.05544
Mass Density GM/CC 0.690 0.772
0.699 0.6904
0.00068789
0.6972487
0.0005732
0.99876
0.998767
Average Molecular Weight
217.32
226.4
224.70
217.32
18.01528 224.6574 20.73324 18.0152
18.01528
49
Unit Operating Conditions B1 - Distillation Column 1
Top Stage Temperature 46.4611115 C Heat duty -136807.949 cal/sec Distillate rate 0.445020567 kmol/hr Reflux rate 61.9355354 kmol/hr Reflux ratio 139.174546
Bottom Stage Temperature 67.071943 C Heat duty 67410.9479 cal/sec Bottoms rate 28.4647816 kmol/hr Boilup rate 28.4647816 kmol/hr Boilup ratio 1
B2 - Inline Mixer 1 Outlet Temp 63.701 C
Pressure 1 bar Vapour Fraction 0
1st Liquid/Total Liquid 1
B3 - Steam Stripper Top Stage
Temperature 140.115998 C Heat duty 0 cal/sec Distillate rate 184.9148 kmol/hr Reflux rate 67.2220849 kmol/hr Reflux ratio 0.363530042
Bottom Stage Temperature 131.061646 C Heat duty 0 cal/sec Bottoms rate 65.0489101 kmol/hr Boilup rate 180.956318 kmol/hr Boilup ratio 2.78185012
B4 - Extraction Column Top stage Temp 23.658 C
Top Stage 1st Liquid Flow 32.663 kmol/h Top Stage 2nd Liquid Flow 69.964 kmol/h
Pressure Change 1 bar NPSH Available 5.0137 m-kgf/kg
Head Developed 14.767 m-kgf/kg Pump Efficiency 0.5314
Net Work Required 1.1509 kW
B12 - Inline Mixer 2 Outlet Temp 128.357 C
Pressure 1 bar Vapour Fraction 0
1st Liquid/Total Liquid 1
51
B13 - Heat Exchanger 2 Outlet Temp 20 C
Pressure 1 bar Vapour Fraction 0
Heat Duty -257222 cal/sec Net Duty -257222 cal/sec
1st Liquid/Total Liquid 1
Economic Tables and Investment Calculations
Table 33: Summary of Factors affecting NPV
Category Grassroots Capital Operating Costs Income Distillation Column Sizing Oil Cost Product Sale Price Steam Stripper Sizing Water Cost Interest Rate Pump Power Rating Maintenance Mixer Power Rating Spare Parts Heat Exchanger Surface Area Supervisors Lab Work License Fees Research and Development Feedstock Storage Product Storage
52
Table 34: NPV and Pay off time calculation summary