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Design and Control of GTBE Process for biofuels.
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Industrial & Engineering Chemistry Research is published by the American ChemicalSociety. 1155 Sixteenth Street N.W., Washington, DC 20036Published by American Chemical Society. Copyright American Chemical Society.However, no copyright claim is made to original U.S. Government works, or worksproduced by employees of any Commonwealth realm Crown government in the courseof their duties.
ArticleDesign and Control of Glycerol Tertiary Butyl EthersProcess for the Utilization of a Renewable Resource
JianKai Cheng, Cheng-Lin Lee, Yong-Tang Jhuang, Jeffrey Daniel Ward, and I-Lung ChienInd. Eng. Chem. Res., Just Accepted Manuscript DOI: 10.1021/ie2010516 Publication Date (Web): 29 Sep 2011
Downloaded from http://pubs.acs.org on October 6, 2011
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1
Paper submitted for publication in Industrial & Engineering Chemistry Research
Design and Control of Glycerol Tertiary Butyl Ethers
Process for the Utilization of a Renewable Resource
Jian Kai Cheng1, Cheng-Lin Lee
2, Yong-Tang Jhuang
2, Jeffrey D. Ward
1, and
I-Lung Chien1*
1 Department of Chemical Engineering,
National Taiwan University,
Taipei 10617, Taiwan
2 Department of Chemical Engineering,
National Taiwan University of Science and Technology,
Taipei 10607, Taiwan
Second Revision: September 20, 2011
* Corresponding author. I-Lung Chien, Tel: +886-3-3366-3063; Fax: +886-2-2362-3040; E-mail:
ilungchien@ntu.edu.tw
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Abstract
In this paper, the design and control of an improved process for the manufacture
of a fuel additive (glycerol tertiary butyl ethers, GTBE) from glycerol and isobutylene
is developed. The improved process redirects one recycle stream, uses a stripping
column instead of a flash tank to recover isobutylene, and uses a rectifying column
instead of a distillation column to purify the product. Economic analysis shows that
the improved process has a 22% lower total annual cost (TAC) than the best known
process published in the literature. Significant increases in the selectivity of the
overall process from 84.7% to 99.3% can also be realized by comparing the optimized
improved design versus the original design. Dynamic simulations were also
conducted and indicated that stringent product specification can be met with a simple
decentralized feedback control structure despite impurities in the feed streams and
also changes in the throughput.
Keywords: GTBE, Glycerol, Fuel additive, Extraction, optimal design, process
dynamics and control
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1. Introduction
Due to increased interest in biodiesel production, the development of new products
and processes that utilize glycerol (a by-product in biodiesel manufacture) has become
important. One possibility is etherification of glycerol with isobutene to produce
glycerol tertiary butyl ethers (GTBEs). The products of the etherification can be a
mono-ether (MTBG), a di-ether (DTBG), or a tri-ether(TTBG), where the higher
ethers (di- and tri-, denoted h-GTBEs) can be used as a diesel additive1 or as an octane
booster for gasoline.2 Due to the two hydroxyl groups, the mono-ether is less soluble in
hydrocarbons.
The glycerol etherification reactions usually take place in the liquid phase at
temperatures between 60 and 100 oC so that the operating pressure should be between
15-20 bar.3 The kinetics of the reactions catalyzed by different acid catalysts has been
investigated by several authors. Among homogeneous catalysts, p-toluenesulfonic acid
gave the best performance.4-7 Among heterogeneous catalysts, Amberlyst 15 gave the
best performance.7 Behr and Obendorf2 also reported that the dimerization of isobutene
to form trimethylpenenes does not occur in the presence of p-toluenesulfonic acid while
about 2% yield of trimethylpenenes is produced in the presence of Amberlyst 15.
Trimethylpenenes are undesirable by-products because they form deposits in the engine
during combustion.2, 8
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For process design, dealing with undesired mono-ether (MTBG) is a critical issue
because of its less solubility in hydrocarbons (e.g. gasoline). Up to the present time,
three different process design alternatives for the manufacture of h-GTBEs have been
proposed in the literatures. They are: the ARCO process,9 the Behr and Obendorf
process,2 and the Di Serio, et al. process.
3 In the ARCO process,
9 a decanter is placed
after the reactor so that unconverted glycerol, p-toluenesulfonic acid and MTBG can
then be recovered in the heavy phase and then recycled back to the reactor. The light
phase is fed to a stripping column, followed by an extraction column (use water as
solvent) for further separation. In the Behr and Obendorf process,2 an extraction column
is placed after a reactor and glycerol feed is used as a solvent to extract unconverted
glycerol, p-toluenesulfonic acid and MTBG. The extract stream is recycled to the reactor
while the raffinate stream is fed to a flash tank, followed by a vacuum column for
further separation. Instead of reducing MTBG content from the product, in the Di Serio,
et al. process,3
Free fatty acid ester (FAME) is used as the solvent to extract GBTEs
(including MTBG, DTBG and TTBG) to solve the problem of the low solubility of
MTBG in fuel. A series of extraction steps in this process were proposed. More detailed
description about the above three processes are shown in Section 1 of the web-published
data (see Appendix A).
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In comparison of the above three processes, the Behr and Obendorf process,2
although it requires vacuum operation in a distillation column, is the most simple and
complete design for obtaining higher ethers satisfying product specification. However, it
only provided a conceptual design flowsheet without detailed information on each
stream data. Furthermore, the dynamics and control strategy of this flowsheet were not
mentioned in this paper.
In this work, the conceptual design of the Behr and Obendorf process is further
investigated to establish an optimized design flowsheet. Then, this optimized flowsheet
is compared with a proposed alternative flowsheet which turns out to be more
economical and also improve the overall selectivity of the process. The study of this
improved flowsheet is further extended to propose an overall control strategy to
properly reject various disturbances from the feed stream.
2. Kinetic and Thermodynamic Models
2.1 Kinetic Model
Etherification of glycerol consists of three serial reversible reactions. Glycerol (GL)
reacts with isobutene (IB) stepwise to form mono-tert-butyl ether of glycerol (MTBG),
di-tert-butyl ether of glycerol (DTBG) and tri-tert-butyl ether of glycerol (TTBG). The
reactions are:
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1
1
2
2
3
3
GL IB MTBG
MTBG IB DTBG
DTBG IB TTBG
+
+
+
k
k
k
k
k
k
(1)
The kinetics of the above reversible reactions are described using a power law
model on the basis of the overall molar concentration of component i (Ci) with the
following reaction rate expressions:
GL1 GL IB 1 MTBG
MTBG1 GL IB 1 MTBG 2 MTBG IB 2 DTBG
DTBG2 MTBG IB 2 DTBG 3 DTBG IB 3 TTBG
TTBG3 DTBG IB 3 TTBG
IB1 GL IB 1 MTBG 2 MTBG IB 2 DTBG 3 DTBG
= +
= +
= +
=
= + +
dCk C C k C
dt
dCk C C k C k C C k C
dt
dCk C C k C k C C k C
dt
dCk C C k C
dt
dCk C C k C k C C k C k C
dtIB 3 TTBG+C k C
(2)
Model parameters were taken from Behr and Obendorf2 for glycerol etherification
catalyzed by p-toluenesulfonic acid (pTS). The kinetic model parameters with Arrehnius
form can be seen in Table 1.
2.2 Thermodynamic Model
To account for non-ideal liquid-liquid equilibrium (LLE) and possible
vapor-liquid-liquid equilibrium (VLLE) for this system, the NRTL10
model was used to
calculate the activity coefficients. In the reaction system, there are five components (GL,
IB, MTBG, DTBG and TTBG). Besides these components, water and 1-butene are also
considered as inert components of impure feeds for the later dynamics and control study.
The binary parameters for the IB-GL-MTBG-h-GTBEs system were taken from Behr
and Obendorf.2 The binary parameter of GL-water was available in the Aspen Plus data
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bank. Other thermodynamic parameters were estimated using the UNIFAC method in
Aspen Plus. The NRTL model parameters are listed in Table 2
With the thermodynamic models and the available corresponding binary
parameters, the phase behavior can be predicted using Aspen Plus. Fig. 1 shows
combined ternary liquid-liquid equilibrium diagrams for the IB-GL-MTBG-h-GTBEs
system, in which the predicted results of liquid-liquid equilibrium of the
GL-MTBG-h-GTBEs ternary system gives a good agreement with the experimental
data by Behr and Obendorf.2
A large liquid-liquid equilibrium envelope between GL-IB-h-GTBEs,
GL-IB-MTBG and GL-MTBG-h-GTBEs systems at P=20 bar and T=90 oC are
exhibited, thus, extraction can be used for the separation. Table 3 shows the boiling
point temperatures for pure components and azeotropic temperatures at P=1bar and
P=0.005bar. MTBG and GL are the heavier components and ternary azeotrope of
GL-DTBG-TTBG and a binary azeotrope of GL-TTBG disappear when the pressure
drops below 0.005bar. Unfortunately, there is no experimental azeotropic data available
in the literature. In spite of this, higher boiling points of system components suggested
that a low pressure distillation column would be more preferable.
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3. Steady State Design
As mentioned previously, the Behr and Obendorf process although requiring
vacuum operation in a distillation column, is the most simplified and complete design
for obtaining higher ethers satisfying product specification. In this section, this
conceptual design is further investigated to establish an optimized design flowsheet.
After that, an alternative improved design is proposed.
Fig. 2A shows the optimized Behr and Obendorf process,2 which includes two
CSTRs in series followed by an extraction column, a flash tank and a vacuum
distillation column. There are three recycle streams in the design flowsheet. Glycerol
feed is introduced into the extraction column to extract mono-ether and pTS catalyst
from the reactor effluent and this mixture is then recycled back to the reaction section.
Excess isobutene is recycled back to the reaction section from the vapor stream of the
flash tank. The bottom stream of the distillation column containing materials not
satisfying product specification (mainly GL and MTBG) is also recycled back to the
reaction section. In our preliminary analysis, we observe that there is still some
isobutene in the product stream, which means that the flash tank cannot recover all of
the isobutene. Therefore, a partial condenser is required to remove additional light
isobutene from the product stream. Isobutene must be removed from the product
because it may form oligomers which can form deposits in the engine during
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combustion.2, 8
3.1 Design Procedure
The optimal design flowsheet is determined by minimizing the total annual cost
(TAC):
capital costTAC operating cost
payback period= + (3)
Here, a payback period of 8 years is used. The operating cost includes the costs of steam,
cooling water and electricity. The capital cost includes the costs of reactors, columns,
trays, vacuum system, compressor and heat exchangers. The method of determining the
total annual cost follows the procedure in Seider et al. 11
The cost models for TAC
calculation are shown in Section 2 of the web-published data (see Appendix A).
To determine the product specification, it is assumed that h-GTBEs products are
used to make a fuel by blending 5 wt% h-GTBEs and 95 wt% biodiesel containing 0.01
wt% glycerol, and that the glycerol content in the blended fuel has to meet ASTM 6571
(< 0.02 wt%). Therefore, the glycerol content in the h-GTBEs product must be less than
0.2 wt%.
3.2 Behr and Obendorf Process
In the conceptual design flowsheet of the Behr and Obendorf process, the operating
condition of each unit is listed in Table 4. Design variables to be determined are shown
in italics in Fig. 2A. They are: feed ratio of fresh glycerol to fresh isobutene (FR),
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residence time of the reactors (R-1 and R-2), number of stages in the extraction column
(NEC), number of stages in the vacuum column (NVC), feed location in the vacuum
column (NF,VC) and h-GTBEs composition in the bottoms of the vacuum column
(xVC,hGTBEs). A systematic, sequential design procedure is devised to generate a
near-optimal flowsheet by varying all the design variables identified above, where
variables at each loop are iterated until an optimal design is determined. Detailed
description of the optimization procedure is listed in Section 3 of the web-published
data (see Appendix A).
All the simulations are carried out in Aspen Plus. In the steady-state simulation, the
pressure of the distillation column is assumed with no pressure drop between trays.
Fixed column pressure makes the convergence of simulation easier and gives no
significant effect in TAC. This constraint is released while exporting to the dynamic
simulation with the commercial software automatically calculates the tray pressure
drop.
FR=2.2 (FIB=11.0 kmol/hr and FGL=5 kmol/hr) is considered first as an
illustration example for the following figures displaying how to determine the values of
the design variables. Fig. 3A indicates that there is a minimum feasible value for R-2
when R-1 is specified. The near-optimal residence time of two reactors occurs when
R-1=4 hr and R-2= 0.5 hr. Fig. 3B and Fig. 3C show how the number of stages in the
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extraction column (NEC), the number of stages in the vacuum column (NVC) and the feed
location of the vacuum column (NF,VC) affecting the TAC. Fig. 3D shows a minimum in
TAC occurs when the h-GTBEs composition of the bottoms in the vacuum column
(xVC,hGTBEs) takes the value of 0.006. As the h-GTBEs content increases, the TAC of the
vacuum column decreases, but the TAC of the reactors increases because the recycle
flow rate from the vacuum column (RVC) increases (see Fig. 4). The stream table of the
near-optimal design flowsheet is listed in Fig. 2A.
3.3 Improved Process
The design of the Behr and Obendorf process indicates that the vacuum column is
the most expensive part of the process, and the product stream still contains a small
amount of isobutene (2.2 mol%). Additional isobutene is also lost in the light waste
stream. Although the light waste stream can be recycled to the reactor, an additional
recycle cost (mainly compressor cost) would be required. Note that in this case the
pressure of the light waste stream must be increased from 0.005 bar to 20 bar to be able
to recycle back to the reaction section.
This shortcoming suggests that the process might be improved by further
modification of the IB recovery column and the h-GTBEs purification column.
Therefore, an alternative process is proposed where a multi-stage stripping column is
used in place of the flash tank as in the ARCO process9 to recovery the IB completely,
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and a rectifying vacuum column (without a reboiler) is used to purify the products.
Because a reboiler is not installed in the rectifying vacuum column, more h-GTBEs
product goes down the column. To avoid the reverse reaction, it might be a better choice
to recycle the bottoms of the vacuum column to the extraction column. In this
configuration, the unconverted glycerol and undesired MTBG are extracted first and
then recycled to the reactor (Fig. 2B). In the following, all design variables of the
improved process are varied in a similar sequential iterative manner to generate a
near-optimal design flowsheet.
Fig. 5A indicates that one CSTR (R-1=3.0 hr and R-2=0 hr) gives a smaller TAC as
compared to two CSTRs. Fig. 5B ~ Fig. 5D show the effect of number of stages in the
extraction column, stripping column, and vacuum column (NEC, NSC and NVC
respectively) on the TAC. Fig. 5E shows that increasing the recycle flow rate from the
vacuum column (RVC) leads to increases in the reaction cost and the separation cost.
Consequently, smaller RVC gives a lower TAC, and in this case a minimum value of RVC
is taken at 0.5 kmol/hr. The stream table of the near-optimal design flowsheet for the
improved process is listed in Fig. 2B.
3.4 Comparison
Comparing the stream tables in Figs. 2A and 2B, both final products can meet
glycerol impurity specification of 0.2 wt%. However, the selectivity to produce
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h-GTBEs in the final product of the improved process is much higher than that of the
Behr and Obendorf process. From Fig. 2, the selectivity of the improved process is at
99.3% while that of the Behr and Obendorf process is at 84.7%. Note that the selectivity
is defined as the ratio of the desired product (h-GTBEs) formed (in moles) to the overall
GTBEs in the products stream (in moles) through the overall process. In the improved
process, the isobutene is almost completely consumed (See Fig 2B), which means more
GTBEs are produced. In addition, recycling the bottoms of the vacuum column to the
extraction column leads to less reverse reaction to produce MTBG from h-GTBE.
A comparison of the TAC between the Behr and Obendorf process and the
improved process is shown in Fig. 6. The results show that using a stripping column to
recover the isobutene is more expensive than using a flash tank. However the isobutene
can be almost completely removed in the stripping column so that there is negligible
amount of isobutene in the product stream (See Fig. 2B). Furthermore, using a
rectifying column without a reboiler dramatically reduces the cost of the purification of
the h-GTBEs product. Redirect the recycle stream of the bottoms of vacuum column to
the extraction column (instead of to the reaction section) makes less h-GTBEs product
going back to the reaction section, thus suppress the reverse reaction. Therefore, one
CSTR is sufficient to complete the reaction for the optimized improved process whereas
two CSTRs are required for the optimized Behr and Obendorf process. This, in terms,
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reduces the cost of the reactor system. Overall, the optimized improved process is 22%
less expensive than the Behr and Obendorf process.
The effect of FR is also considered. For the Behr and Obendorf process, Fig. 7A
reveals that the TAC increases as FR decreases. This is because the TAC of the
plant-wide process is dominated by the cost of the vacuum column. A decreasing of FR
means less isobutene is fed to the reactors and consequently the concentration of
product (h-GTBEs) in the feed to the vacuum column is decreased, which leads to
higher cost of the vacuum column. For the improved process, Fig. 7B indicates that the
TAC decreases as FR decreases. This is because the stripping column supplies a larger
recycle flow rate of the isobutene (Fig. 2B), so the excess isobutene is still enough even
if FR decreases. Furthermore, the larger FR leads to larger flow rates of internal
recycles, which increases the cost of all units. The design parameters of the near-optimal
design for the two processes and the corresponding costs of all units are summarized in
Tables 5 and 6.
In the comparison of the two processes at FR of 2.1 or 2.2, Figure 7 shows that the
improved process all gives significantly TAC savings. For the Behr and Obendorf
process at FR=2.3, although gives better TAC than that of the improved process at the
same FR, however, the cost is still at least 19% higher than that of the improved process
at FR=2.1 or 2.2. Notice also that the TAC calculation does not includes the cost
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associated with the feed stream. Thus, if this additional cost of the feed stream with
higher FR is included, more benefit can be realized for the improved process.
4. Process Dynamics and Control
In this section, the dynamics and control of the improved process is investigated by
introducing feed flow and composition disturbances. The case of FR=2.2 is studied. The
throughput changes of the process can be achieved with the feed flow changes. For the
feed composition disturbances, impure feeds conditions are: the glycerol contains
2.5wt% water and the isobutene feed contains 2.5wt% 1-butene. In order to purge out
these impurities to prevent accumulation in the system, a minor modification is made to
the plant-wide process. The extract flow from the extraction column is fed to a new
flash tank (F-1 in Fig. 8) where light impurities are withdrawn as a vapor stream before
the remaining liquid is fed to the reactor. The operating condition of the new flash tank
(F-1) is set at 1 bar and 130 oC. Although a change of an additional unit is made to the
process, further process optimization of this modified design is not considered because
the cost of the flash tank is much less than other columns. Also, including this flash tank
in the TAC calculation will just add a constant in the overall value of TAC, thus will not
affect the design variables of the flowsheet.
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Next, a control structure is developed for the improved process by following these
principles, as shown in Fig. 8.
(1) The glycerol feed flow is selected as the throughput manipulator and there are 15
manipulated variables remaining.
(2) In the entire process, eight inventory control loops are required, which include
control of 5 levels and 3 vessel pressures. Basic inventory and related loops are
arranged as follows: Top pressures of flash tank, stripping column and vacuum
column are controlled by manipulating vapor flow rate, compressor work and
condenser duties respectively. Liquid levels are controlled by manipulating vessel
outlet flow rates.
(3) After selecting the inventory controls, the remaining 7 manipulated variables (heat
duties of reactor, two heat exchangers, and the flash tank, reboiler duty of stripping
column, reflux of vacuum column, isobutene feed flow rate) are used for quality
control, which are determined as below:
i. Control the temperature of the reactor (R-1) to 90 oC by manipulating the heat
duty.
ii. Control the temperatures of the heat exchangers (HX-1 and HX-2) to 40 oC
and 90 oC respectively by manipulating their heat duties.
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iii. Control the temperatures of the flash tank (F-1) to 130 oC by manipulating its
heat duty.
iv. Control the tray temperature of the stripping column by manipulating the
reboiler duty to maintain the top composition. Sensitivity analysis was
performed for 0.01% variations of the reboiler duty showing that the first
tray is most sensitive.
v. Fix the reflux ratio in the vacuum column. Other alternative configurations
(fix reflux flow rate or fix ratio of reflux to feed) are also considered, however,
there is no significant difference in the control performance among them, as
shown in Section 4 of the web-published data (see Appendix A).
vi. Fix the ratio of the mixing flow rate of the fresh isobutene and the recycle
stream from the stripping column (primarily isobutene) to the liquid stream
from the flash tank (mostly glycerol) by adjusting the fresh isobutene feed
flow rate. In this manner, stoichiometric balance into the reactor is
maintained.
After the decentralized control structure was designed, dynamic simulations were
performed using Aspen Plus DynamicsTM
. A third-order 0.5 minute time lag was
assumed for temperature measurement.12
Flow, pressure, and temperature were
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controlled using proportional-integral controllers while proportional-only control was
used for liquid level with Kc=2. Relay feedback tests13
were performed on the
temperature loops to find the ultimate gains (Ku) and ultimate period (Pu) of each
temperature and ratio control loop. Modified Tyreus-Luyben tuning relations (Kc = Ku/3
and I =2 Pu) were used to generate initial controller parameters. In order to obtain an
acceptable damping, further detuning from the initial settings was required for some of
the loops. Controller settings for main control loops are summarized in Table 7.
The plant-wide control is tested for feed flow and composition disturbances. Fig. 9
shows that fast responses can be obtained for the disturbance sequence: increasing
glycerol fresh feed rate from 5 to 6 kmol/hr at t= 5 hr, decreasing glycerol fresh feed
flow rate from 6 to 4 kmol/hr at t=80 hr, returning the original production rate at t=160
hr and finally introducing impurity in both feeds (97.5wt% glycerol/2.5wt% water and
97.5wt% isobutene/2.5wt% 1-butene) at t=240 hr.
All the temperature control loops settle quickly in less than 5 hours. The
compositions of the final product stream usually settle slower than the temperature
control loops because of recycle loops in the overall process to defer all components to
reach new steady-state. However, as can be seen in Fig. 9 all compositions of the final
product stream also settle quite quickly (less than 40 hours). More importantly, all
responses under various feed disturbances give very small offsets in the product purity.
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As expected, the flow rate of the final product increases/decreases accordingly to the
feed flow changes.
For the feed impurity disturbances, the addition of the new flash tank successfully
purges out the impurities in the two feed streams. The glycerol impurity in the final
product actually decreases to a smaller amount. The reduction of the glycerol content in
the h-GTBEs product stream is because the water in the glycerol feed helps to wash
down glycerol in the extraction column. Overall, reasonable control performance can be
obtained using a simple control structure for production rate and feed impurity
concentration variations.
5. Conclusion
An improved design for the manufacture of h-GTBE has been presented.
Economic analysis shows that the improved design has a 22% lower total cost than the
optimized design reported in the literature. Significant increases of the selectivity of the
overall process from 84.7% to 99.3% can also be realized. The plant-wide control
structure of the improved design is also developed. Dynamic simulations indicate that
stringent product specification can be met with this simple control structure despite
impurities in the feed streams and also changes in the throughput.
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Appendix A. Supplementary data
Supplementary data associated with this article can be found, in the online version,
at doi: xxx.
Acknowledgement
The research funding from the National Science Council and from Ministry of
Economic Affair of the R. O. C. are greatly appreciated.
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References:
1. Jaecker-Voirol, A.; Durand, I.; Hillion, G.; Delfort, B.; Montagne, X. Glycerin for
New Biodiesel Formulation. Oil Gas Sci. Technol. 2008, 63.
2. Behr, A.; Obendorf, L. Development of a Process for the Acid-Catalyzed
Etherification of Glycerine and Isobutene Forming Glycerine Tertiary Butyl Ethers.
Eng. Life Sci. 2002, 2.
3. Di Serio, M.; Casale, L.; Tesser, R.; Santacesaria, E. New Process for the
Production of Glycerol Tert-Butyl Ethers. Energy & Fuels 2010, 24.
4. Behr, A.; Obendorf, L. Process Development for Acid-Catalysed Etherification of
Glycerol with Isobutene to Form Glycerol Tertiary Butyl Ethers. Chemie
Ingenieur Technik 2001, 73.
5. Klepacova, K.; Mravec, D.; Bajus, M. Tert-Butylation of Glycerol Catalysed by
Ion-Exchange Resins. Appl. Catal. A-Gen. 2005, 294.
6. Karinen, R. S.; Krause, A. O. I. New Biocomponents from Glycerol. Appl. Catal.
A-Gen. 2006, 306.
7. Klepacova, K.; Mravec, D.; Kaszonyi, A.; Bajus, M. Etherification of Glycerol
and Ethylene Glycol by Isobutylene. Appl. Catal. A-Gen. 2007, 328.
8. Melero, J. A.; Vicente, G.; Morales, G.; Paniagua, M.; Moreno, J. M.; Roldn, R.;
Ezquerro, A.; Prez, C. Acid-Catalyzed Etherification of Bio-Glycerol and
Isobutylene over Sulfonic Mesostructured Silicas. Appl. Catal. A-Gen. 2008, 346.
9. Gupta, V. P. Glycerine Ditertiary Butyl Ether Preparation. U.S. Patent 5,476,971,
1995.
10. Renon, H.; Prausnit.Jm Local Compositions in Thermodynamic Excess Functions
for Liquid Mixtures. AIChE J. 1968, 14.
11. Seider, W. D.; Seader, J. D.; Lewin, D. R.; Widagdo, S.; Product and Process
Design Principles : Synthesis, Analysis, and Evaluation 3rd ed.; John Wiley:
Hoboken, NJ, 2009.
12. Luyben, W. L.; Tyreus, B. D.; Luyben, M. L.; Plantwide Process Control;
McGraw-Hill: New York, 1998.
13. Shen, S. H.; Yu, C. C. Use of Relay-Feedback Test for Automatic Tuning of
Multivariable Systems. AIChE J. 1994, 40.
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Table 1 Kinetic Parameters of the etherification reaction (Behr and Obendorf2)
pre-exponential factor Activation Energy
k0,1 [sec-1kmol-1] 5.07106 E1 [kJ/kmol] 7.404104
k0,-1 [sec-1] 6.151011 E-1 [kJ/kmol] 1.118105
k0,2 [sec-1kmol-1] 2.83109 E2 [kJ/kmol] 9.280104
k0,-2 [sec-1] 1.421013 E-2 [kJ/kmol] 1.181105
k0,3 [sec-1kmol-1] 3.77108 E3 [kJ/kmol] 9.256104
k0,-3 [sec-1] 1.061014 E-3 [kJ/kmol] 1.251105
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Table 2 NRTL model Parameters of the glycerol etherification system
Comp. i GL GL GL GL MTBG MTBG MTBG
Comp. j MTBG DTBG TTBG IB DTBG TTBG IB
aij 0 0 0 0 0 0 0
aji 0 0 0 0 0 0 0
bij 207.34 1573.3 1573.3 721.75 -630.83 -630.83 -310.24
bji 79.22 528.53 528.53 937.02 680.4 680.4 1229.9
cij 0.2 0.2 0.2 0.2 0.2 0.2 0.2
Source BO BO BO BO BO BO BO
Comp. i DTBG TTBG DTBG GL GL IB IB
Comp. j IB IB Water TTBG 1-BUTENE 1-BUTENE WATER
aij 0 0 -0.732 0 0 0 0
aji 0 0 -1.252 0 0 0 0
bij -742.47 -742.47 170.92 114.94 658.75 107.53 830.03
bji -465.14 -465.14 272.61 56.71 2363.90 -93.25 1588.00
cij 0.2 0.2 .03 0.3 0.3 0.3 0.3
Source BO BO Aspen UNIFAC UNIFAC UNIFAC UNIFAC
Comp. i MTBG MTBG DTBG DTBG TTBG TTBG 1-BUTENE
Comp. j 1-BUTENE WATER 1-BUTENE WATER 1-BUTENE WATER WATER
aij 0 0 0 0 0 0 0 aji 0 0 0 0 0 0 0
bij -39.33 -526.56 -337.01 -342.76 1138.84 219.95 795.81
bji 1352.67 1947.31 802.78 2861.95 -660.15 4153.94 1629.56
cij 0.3 0.3 0.3 0.3 0.3 0.3 0.3
Source UNIFAC UNIFAC UNIFAC UNIFAC UNIFAC UNIFAC UNIFAC
UNIFAC: Predicted by UNIFAC Method in Aspen Plus
Aspen: Built in the databank of Aspen Plus
BO: Binary parameters are given by Behr and Obendorf2
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Table 3 Boiling point ranking for pure components and azeotropes at P=1bar and
P=0.005bar
P=1bar
Component / azeotrope Composition
(mole fraction) Temperature (C)
IB - -7.42
GL/DTBG/TTBG 0.198/0.600/0.202 233.65
GL/DTBG 0.205/0.795 234.51
DTBG/TTBG 0.768/0.232 239.00
DTBG - 239.96
GL/TTBG 0.266/0.734 242.25
TTBG - 252.45
MTBG - 264.22
GL - 287.21
P=0.005bar
Component / azeotrope Composition
(mole fraction) Temperature (C)
IB - -92.83
DTBG/TTBG 0.163/0.837 86.63
TTBG - 87.10
DTBG - 96.38
GL/DTBG 0.002/0.998 96.38
MTBG - 132.11
GL - 147.53
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Table 4 Operating Conditions for Process Units
Units Pressure (bar) Temperature (oC)
R-1 20 90
R-2 20 90
HX-1 20 40
HX-2 20 90
EC-1 20 -
F-1 1 100
VC-1 0.005 -
Comp-1 5 -
SC-1 1 -
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Table 5 Steady-state design of the Behr and Obendorf process
Cases
Equipments FR=2.1 FR=2.2 FR=2.3
1st CSTR (R-1)
Residence Time [hr] 1 4 2
Volume [m3] 4.90 7.91 4.22
Heat Transfer Area [ft2] 6.06 3.66 6.37
TAC [$1000/yr] 27.59 49.28 35.07
Capital Cost [$1000] 211.25 393.73 279.72
Operating Cost [$1000/yr] 1.18 0.06 0.10
2nd CSTR (R-2)
Residence Time [hr] 1.5 0.5 2
Volume [m3] 6.83 0.97 3.96
Heat Transfer Area [ft2] 7.92 1.87 5.4
TAC [$1000/yr] 31.46 16.81 33.89
Capital Cost [$1000] 250.71 134.22 270.40
Operating Cost [$1000/yr] 0.13 0.03 0.09
Extractor (EC-1)
Total Number of Tray 3 3 3
Number of Feed Tray 1 and 3 1 and 3 1 and 3
TAC [$1000/yr] 5.41 4.96 5.03
Capital Cost [$1000] 43.35 39.70 40.23
Flash tank(F-1)
Heating Duty [kW] 43.24 45.24 52.87
TAC [$1000/yr] 11.56 11.88 12.93
Capital Cost [$1000] 68.83 70.29 74.48
Operating Cost [$1000/yr] 2.96 3.10 3.62
Compressor (Comp-1)
Electrical input[kW] 0.88 0.81 0.86
TAC [$1000/yr] 3.83 3.58 3.77
Capital Cost [$1000] 27.27 25.57 26.88
Operating Cost [$1000/yr] 0.42 0.38 0.41
Vacuum Column (VC-1)
Total Number of Tray 19 19 19
Number of Feed Tray 18 18 18
Diameter of Column [m] 2.28 2.11 1.91
Duty of Condenser [kW] 475.70 393.92 305.37
Duty of Reboiler [kW] 459.91 378.18 291.91
TAC [$1000/yr] 242.86 219.18 194.49
Capital Cost [$1000/yr] 1447.13 1345.62 1241.03
Operating Cost [$1000/yr] 61.98 50.98 39.36
Vacuum system
TAC [$1000/yr] 3.09 3.07 3.05
Capital Cost [$1000] 14.59 14.55 14.50
Operating Cost [$1000/yr] 1.26 1.25 1.24
1st Heat exchanger (HX-1)
Heating Duty [kW] -67.17 -56.02 -57.42
TAC [$1000/yr] 2.01 1.90 1.92
Capital Cost [$1000] 13.17 12.80 12.85
Operating Cost [$1000/yr] 0.36 0.30 0.31
2nd Heat exchanger (HX-2)
Heating Duty [kW] -2.78 -2.54 -2.72
TAC [$1000/yr] 0.80 0.79 0.80
Capital Cost [$1000] 6.29 6.20 6.27
Operating Cost [$1000/yr] 0.015 0.014 0.015
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Table 6 Steady-state design of the improved process
Cases
Equipments Cases FR=2.1 FR=2.2 FR=2.3
1st CSTR (R-1)
Residence Time [hr] 1.5 3 6
Volume [m3] 4.71 7.60 16.02
Heat Transfer Area [ft2] 7.84 0.11 2.53
TAC [$1000/yr] 38.66 47.65 74.28
Capital Cost [$1000] 297.05 380.99 593.93
Operating Cost [$1000/yr] 1.53 0.02 0.04
Extractor (EC-1)
Total Number of Tray 3 5 6
Number of Feed Tray 1 and 3 1 and 5 1 and 6
TAC [$1000/yr] 6.54 7.93 9.20
Capital Cost [$1000] 52.34 63.45 73.57
Stripping column (SC-1)
Total Number of Tray 5 5 5
Number of Feed Tray 1 1 1
Diameter of Column [m] 0.45 0.47 0.52
Duty of Condenser [kW] - - -
Duty of Reboiler [kW] 250.44 261.26 299.12
TAC [$1000/yr] 98.20 100.61 122.98
Capital Cost [$1000] 391.32 393.57 398.32
Operating Cost [$1000/yr] 49.28 51.41 73.19
Compressor (Comp-1)
Electrical input[kW] 18.73 20.49 27.45
TAC [$1000/yr] 40.05 42.97 54.11
Capital Cost [$1000] 249.49 266.25 329.02
Operating Cost [$1000/yr] 8.86 9.69 12.98
Vacuum Column (VC-1)
Total Number of Tray 2 2 2
Number of Feed Tray 2 2 2
Diameter of Column [m] 1.16 1.19 1.22
Duty of Condenser [kW] 101.24 105.10 109.03
Duty of Reboiler [kW] - - -
TAC [$1000/yr] 38.32 37.89 37.62
Capital Cost [$1000] 300.04 296.34 293.91
Operating Cost [$1000/yr] 0.82 0.85 0.88
Vacuum system
TAC [$1000/yr] 2.97 2.97 2.98
Capital Cost [$1000] 14.15 14.26 14.26
Operating Cost [$1000/yr] 1.19 1.19 1.19
1st Heat exchanger (HX-1)
Heating Duty [kW] 89.45 68.10 67.35
TAC [$1000/yr] 2.2 2.01 2.01
Capital Cost [$1000] 13.75 13.17 13.16
Operating Cost [$1000/yr] 0.48 0.37 0.36
2nd Heat exchanger (HX-2)
Heating Duty [kW] 131.83 139.10 177.62
TAC [$1000/yr] 2.26 2.32 2.59
Capital Cost [$1000] 12.44 12.57 13.10
Operating Cost [$1000/yr] 0.71 0.75 0.96
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Table 7 Final controller settings
Controlled
Variables
Manipulated
Variables Kc [-] I [min] PV Range OP Range
TR-1 QR-1 10.27 11.9 40-140 [oC] -116-0 [kW]
RIB,GLa FIB 10.28 1.2 0-16.4 [-] 0-22 [kmol/hr]
T1,SC QR,SC 1.23 7.1 0-373 [oC] 0-557.2 [kW]
TF QR,F 33.0 5.28 0-260 [oC] 0-123.3 [kW]
PF-1 VF-1 2 10 0-2 [bar] 0-2.28 [kmol/hr]
PSC-1 WComp-1 2 10 0-2 [bar] 0-44.4[kW]
PVC-1 QVC-1 2 10 0-0.01 [bar] -222-0 [kW] aRIB,GL: Ratio of the mixing flow rate of the fresh isobutene and the recycle stream from
the stripping column to the liquid stream from the flash tank.
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Figure Captions
Fig. 1. Predicted liquid-liquid equilibrium (in mass fraction) using ASPEN PLUS at
P=20 bar and T=90oC. (Experimental data is given by Behr and Obendorf2)
Fig. 2. The optimized flowsheets: (A) The Behr and Obendorf process and (B) The
improved process.
Fig. 3. Effect of design varaibles on TAC for the Behr and Obendorf process: (A)
residence times in the reactors (R-1 and R-2); (B) Number of trays in the
extraction column (NEC); (C) Number of trays in the vacuum column and
vacuum column feed tray (NVC and NF,VC); (D) h-GTBEs composition of the
bottoms in the vacuum column (xVC,hGTBEs).
Fig. 4. The relationship between the h-GTBEs composition of the bottoms in the
vacuum column (xVC,hGTBEs) and the recycle flow rate from the vacuum column
(RVC).
Fig. 5. Effects of design varaibles to TAC for the improved process: (A) residence times
in the reactors (R-1 and R-2); (B) Number of trays in the extraction column (NEC);
(C) Number of trays in the stripping column (NSC); (D) Number of trays in the
vacuum column (NVC); (E) Recycle flow rate from the vacuum column (RVC).
Fig. 6. Comparison of TAC between the Behr and Obendorf process and the improved
process.
Fig. 7. Effects of FR on TAC: (A) the Behr and Obendorf process and (B) the improved
process.
Fig. 8. Plant-wide control structure of the improved process.
Fig. 9. Dynamic response for the following load sequence: increasing glycerol fresh
feed rate from 5 to 6 kmol/hr at t=5 hr, decreasing glycerol fresh feed flow rate
from 6 to 4 kmol/hr at t=80 hr, returning the original production rate at t=160 hr
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and finally introducing impurity in both feeds (97.5wt% glycerol/2.5wt% water
and 97.5wt% isobutene/2.5wt% 1-butene) at t=240 hr.
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0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9
Ternar
y Map
( Mas s
B asi s
)
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
0.10.20.30.40.50.60.70.80.9
0.9
0.8
0.7
0.6
0.5
0.4
0.3
0.2
0.1
Ternary Map (Mass Basis)
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
0.9
0.8
0.7
0.6
0.5
0.4
0.3
0.2
0.1
Ter nary Map (M as s Bas is )
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1
IBGL
h-GTBEs
MTBG
MTBG MTBG
Experimental data
Calculation data
Fig. 1. Predicted liquid-liquid equilibrium (in mass fraction) using ASPEN PLUS at
P=20 bar and T=90oC. (Experimental data is given by Behr and Obendorf2)
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Fig. 2. The optimized flowsheets: (A) The Behr and Obendorf process and (B) The
improved process.
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0.5 1.0 1.5 2.0 2.5 3.0
310
315
320
325
330
335
R-1=4.5
R-1=4.0
R-1=3.5
R-1=3.0
NEC=3, N
VC=20, N
F,VC=19, x
VC,hGTBEs=0.006
TAC ($1000/yr)
R-2 (hr)
3 4 5 6 7311
312
313
314
315
316
R-1=4,
R-2=0.5, N
VC=20, N
F,VC=19, x
VC,hGTBEs=0.006
TAC ($1000/yr)
NEC
15 16 17 18 19 20 21
310
320
330
340
350
360 N
VC=18
NVC=19
NVC=20
R-1=4,
R-2=0.5, N
EC=3, x
VC,hGTBEs=0.006
TAC ($1000/yr)
NF,EC
0.003 0.004 0.005 0.006 0.007 0.00850
200
250
300
350
TACOverall
TACVC
TACReactors
R-1=4,
R-2=0.5, N
EC=3, N
VC=20, N
F,VC=19
TAC ($1000/yr)
xVC,hGTBEs
Fig. 3. Effect of design varaibles on TAC for the Behr and Obendorf process: (A)
residence times in the reactors (R-1 and R-2); (B) Number of trays in the extraction
column (NEC); (C) Number of trays in the vacuum column and vacuum column feed tray
(NVC and NF,VC); (D) h-GTBEs composition of the bottoms in the vacuum column
(xVC,hGTBEs).
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0.003 0.004 0.005 0.006 0.007 0.008
1
2
3
4
5
6
7
8
R-1=4,
R-2=0.5, N
EC=3, N
VC=20, N
F,VC=19
RVC (kmol/hr)
xVC,hGTBEs
Fig. 4. The relationship between the h-GTBEs composition of the bottoms in the
vacuum column (xVC,hGTBEs) and the recycle flow rate from the vacuum column (RVC).
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(A) (B)
(C) (D)
(E)
1.0 1.5 2.0 2.5 3.0 3.5 4.0
245
250
255
260
265
R-2=1.0
R-2=0.5
R-2=0
NEC=5, N
SC=5, N
VC=2, R
VC=0.5
TAC ($1000/yr)
R-1 (hr)
2 3 4 5 6
245
250
255
260
R-1=3,
R-2=0, N
SC=5, N
VC=2, R
VC =0.5
TAC ($1000/yr)
NEC
3 4 5 6 7240
260
280
300
R-1=3,
R-2=0, N
EC=5, N
VC=2, R
VC=0.5
TAC ($1000/yr)
NSC
2 3 4 5244
246
248
250
252
254
R-1=3,
R-2=0, N
EC=5, N
SC=5, R
VC=0.5
TAC ($1000/yr)
NVC
0.5 1.0 1.5 2.0455055
200
220
240
260
TACOverall
TACSep
TACRX
R-1=3,
R-2=0, N
EC=5, N
SC=5, N
VC=2
TAC ($1000/yr)
RVC
Fig. 5. Effects of design varaibles to TAC for the improved process: (A) residence times
in the reactors (R-1 and R-2); (B) Number of trays in the extraction column (NEC); (C)
Number of trays in the stripping column (NSC); (D) Number of trays in the vacuum
column (NVC); (E) Recycle flow rate from the vacuum column (RVC).
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0
100
200
300
18%
897%
145%
72%
78%
100%
100%100%
100%
100%
hGTBE Purification System
IB Recovery System
Extraction System
Reactor System
Plant-wide Process
TAC [$1000/yr]
The Berh and Obendorf Process The Alternative Process
Fig. 6. Comparison of TAC between the Behr and Obendorf process and the improved
process.
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(A)
0
50
100
150
200
250
300
350
TAC [$1000/yr]
FR=2.1
FR=2.2 FR=2.3
hGTBE Purification System
IB Recovery System
Extraction System
Reactor System
Plant-wide Process
(B)
0
50
100
150
200
250
300
350
hGTBE Purification System
IB Recovery System
Extraction System
Reactor System
Plant-wide Process
TAC [$1000/yr]
FR=2.1 FR=2.2 FR=2.3
Fig. 7. Effects of FR on TAC: (A) the Behr and Obendorf process and (B) the improved
process.
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GL
h-GTBE
LC
PC
X
FT FT
FCSP
LC
LC LC
PC
TC
TC
TC
FT
IB
LC
PC
TC
TC1
Light
Waste
R-1
F-1
HX-1
EC-1
SC-1 VC-1
HX-2
Comp-1
FC
FT
RC
FCSP
Fig. 8. Plant-wide control structure of the improved process.
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39
050
100
150
200
250
300
350
456
050
100150200250300350
0.004
0.008
0.012
0.016
050
100150200250300350
0.000
0.001
0.002
0.003
0.004
050
100
150
200
250
300350
68
10
12
14
050
100150200250300350
0.0
0.1
0.2
0.3
0.4
050
100150200250300350
0.019
0.020
0.021
0.022
0.023
050
100
150
200
250
300350
0.71
0.72
0.73
0.74
0.75
050
100150200250300350
89.8
89.9
90.0
90.1
90.2
050
100
150200
250300
350
4.0
4.5
5.0
5.5
6.0
050
100
150200
250300
350
184
186
188
190
050
100150200250300350
0.24
0.25
0.26
0.27
0.28
050
100150200250300350
-100
-80
-60
-40
-20
050
100
150
200
250
300
350
200
250
300
350
050
100150200250300350
1.65
1.70
1.75
1.80
050
100
150
200
250
300
350
68
10
12
14
050
100150200250300350
0.65
0.70
0.75
0.80 0
50
100150
200250
300350
0.12
0.13
0.14
0.15
050
100
150
200
250
300
350
8
10
12
14
050
100
150
200
250
300
350
0.40
0.45
0.50
0.55
0.60
0.65 0
50
100
150
200
250
300
0.26
0.28
0.30
0.32
0.34
050
100150200250300350
0.6
0.8
1.0
1.2
1.4
050
100150200250300350
10
12
14
16
18
050
100150200250300350
0.72
0.74
0.76
0.78
0.80
0.82
050
100150200250300350
0.115
0.120
0.125
0.130
0.1350
50
100150200250300350
0.2
0.4
0.6
0.8
1.0
1.2
050
100150200250300350
0.4
0.6
0.8
1.0
050
100150200250300350
0.0
0.1
0.2
0.3
0.4
0.5
050
100150200250300350
0.0
0.1
0.2
0.3
0.4
050
100150200250300350
129
130
131
132
050
100150200250300350
40.0
50.0
60.0
70.0
80.0
F_GL (kmol/hr)
Time (hr)
Raffinate_MTBG (mol/mol)
Time (hr)
Product_GL (kg/kg)
Time (hr)
F_IB (kmol/hr)
Time(hr)
RR (kmol/kmol)
Time (hr)
Raffinate_GL (mol/mol)
Time (hr)
Product_DTBG (kg/kg)
Time (hr)
CSTR1_T (
o
C)
Time (hr)
Product_F (kmol/hr)
Time (hr)
SC_T (
o
C)
Time (hr)
Product_TTBG (kg/kg)
Time (hr)
CSTR1_Q (kw)
Time (hr)
SC_Q (kw)
Time (hr)
Ratio (IB
R
/GL
R
)
Time (hr)
SC_Distillate (kmol/hr)
Time (hr)
Distillate_IB (mol/mol)
Time (hr)
Distillate_DTBG(mol/mol)
Time (hr)
Extract (kmol/hr)
Time (hr)
Extract_GL (mol/mol)
Time (hr)
Extract_MTBG (mol/mol)
Time (hr)
VC_Bottoms (kmol/hr)
Time (hr)
Raffinate (kmol/hr)
Time (hr)
VC_Bottoms_DTBG (mol/mol)
Time (hr)
VC_Bottoms_TTBG (mol/mol)
Time (hr)
Purge_F (kmol/hr)
Time (hr)
Purge_IB (mol/mol)
Time (hr)
Purge_H2O (mol/mol)
Time (hr)
Purge_1-Butene (mol/mol)
Time (hr)
F1_T (
o
C)
Time (hr)
F1_Q (kw)
Time (hr)
Fig. 9. Dynamic response for the following load sequence: increasing glycerol fresh
feed rate from 5 to 6 kmol/hr at t=5 hr, decreasing glycerol fresh feed flow rate from 6
to 4 kmol/hr at t=80 hr, returning the original production rate at t=160 hr and finally
introducing impurity in both feeds (97.5wt% glycerol/2.5wt% water and 97.5wt%
isobutene/2.5wt% 1-butene) at t=240 hr.
Page 39 of 39
ACS Paragon Plus Environment
Industrial & Engineering Chemistry Research
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