Design and Control of GTBE Process for the Utilization of a Renewable Resource IEC Research Oct 2011

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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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:

[email protected]

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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-1] 1.421013 E-2 [kJ/kmol] 1.181105


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


TAC [$1000/yr] 31.46 16.81 33.89

Capital Cost [$1000] 250.71 134.22 270.40


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


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


Vacuum Column (VC-1)


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


Vacuum system

TAC [$1000/yr] 3.09 3.07 3.05

Capital Cost [$1000] 14.59 14.55 14.50


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


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


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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


TAC [$1000/yr] 38.66 47.65 74.28

Capital Cost [$1000] 297.05 380.99 593.93


Extractor (EC-1)


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)




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


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


Vacuum Column (VC-1)




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


Vacuum system

TAC [$1000/yr] 2.97 2.97 2.98

Capital Cost [$1000] 14.15 14.26 14.26


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


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


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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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36

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


IB Recovery System

Extraction System

Reactor System

Plant-wide Process

(B)

0

50

100

150

200

250

300

350


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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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



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Design and Control of GTBE Process for the Utilization of a Renewable Resource IEC Research Oct 2011

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