Fuel Additives

11 FUEL ADDITIVES CHAPTER OUTLINE 11.1 Introduction 400 11.2 Glycerol Etherification With Isobutene 11.2.1 Basis of Design 402 11.2.1.1 Chemical Reaction Network 11.2.1.2 Physical Properties 402 11.2.1.3 Reaction Kinetics 404

400 402

11.2.2 Conceptual Design 406 11.2.3 Process Design 406 11.3 Glycerol Etherification With Tert-butanol 11.3.1 Basis of Design 409 11.3.1.1 Reaction Stoichiometry and Kinetics 11.3.1.2 Physical Properties 411 11.3.1.3 Vapor Liquid Equilibrium 411

11.3.2 ReactoreSeparationeRecycle Process

408 409

413

11.3.2.1 Plant Flowsheet 413 11.3.2.2 Dynamics and Control 416

11.3.3 Reactive Distillation for Glycerol Etherification with Tert-butyl Alcohol 418 11.3.3.1 Reactive Distillation Design 418 11.3.3.2 Dynamics and Control 420 11.3.3.3 Economic Evaluation 422

11.4 Glycerol Ketalization 422 11.4.1 Basis of Design 423 11.4.2 Conceptual Design 424 11.4.3 Design of the Chemical Reactor 425 11.4.4 Design of the Separation Section 426 11.4.5 Plantwide Control 427 11.5 Glycerol Acetalization 430 11.5.1 Basis of Design 431 11.5.2 Chemical Reaction 432 11.5.3 Conceptual Design 433

Applications in Design and Simulation of Sustainable Chemical Processes. https://doi.org/10.1016/B978-0-444-63876-2.00011-5 Copyright © 2019 Elsevier B.V. All rights reserved.

399

400

Chapter 11 FUEL ADDITIVES

11.5.4 Reactor Design 434 11.5.5 Separation Section 435 11.5.6 Process Control 436 11.6 Concluding Remarks 437 References 438

11.1 Introduction Biodiesel fuel is a mixture of fatty acid esters obtained by the transesterification reaction of triglycerides from various natural sources, such as vegetable oils and animal fats, with low alcohols, such as methanol. During the transesterification reaction, glycerol is obtained as a valuable by-product, equivalent to approximately 10% wt. of the total biodiesel production. As a result, large amounts of glycerol become available worldwide. At the end of 2013, the worldwide production of glycerol was about 3 million metric tons, from which more than 60% was because of biodiesel (www.oleoline.com). New applications of glycerol as building block molecule emerged (Rahmat et al., 2010; Tan et al., 2013). A rational way of valorization is converting the glycerol in higher value additives for diesel fuels. This chapter presents several alternatives of converting glycerol to more valuable products. Thus, reaction with isobutene or tert-butanol leads to ethers that have a structure which is similar to other ethers used for long time in industry as additives for gasoline, such as MTBE,
cova
et al., 2003; Behr et al., 2008). ETBE, and TAME (Klepa Reaction of glycerol with an aldehyde (such as formaldehyde, acetaldehyde, n-butanal, up to n-decanal) leads to acetals, which find applications as fuel additives, surfactants, flavors, or disinfectants (Silva et al., 2010; Agirre et al., 2011a, Agirre et al., 2011b). Similarly, glycerol can react with ketones, among which acetone is the most common one (Nanda et al., 2014). The product of this reactiondknown as solketaldreduces the freezing point and enhances the lubricant properties of (bio)diesel, while also having applications in cosmetic and pharmaceutical industry.

11.2 Glycerol Etherification With Isobutene The etherification of glycerol with different alcohols or with alkenes leads to products similar as structure to the well-known MTBE, ETBE, and TAME. In particular, glycerol tert-butyl ethers (GTBE) are good fuel additives (Karinen and Krause, 2006), contributing to reduction of soot and particulate matter,

Chapter 11 FUEL ADDITIVES

which result because of incomplete diesel fuel combustion. This is an issue of major concern, especially in some large cities where the authorities plan to limit or even forbid the access of diesel vehicles. Although today high-efficiency filters are compulsory for new cars and trucks, using appropriate additives allows in situ solution of this problem. Research conducted by scientific organization as TNO in the Netherlands showed that by using 2% w/w GTBE as fuel additive, the PM10 could be reduced by about 30% and PM2.5 by about 80%. The biodiesel production in EU reached 10 Mtpy in 2010, corresponding to 1 Mtpy glycerol. Assuming that 80% is converted into a product with the molecular weight of 211 (corresponding to a product composition of 13% mono-GTBE, 64% di-GTBE, and 22% tri-GTBE), 1.84 Mtpy GTBE would result (Jaecker-Voirol et al., 2008). Employing 2% GTBE additive would ensure the treatment of 92 Mtpy diesel fuel or about 60% of an annual European consumption of 150 Mtpy. Another alternative studied by the French Petroleum Institute is manufacturing a new fuel with the composition 7.5% GTBE and 92.5% FAME (Jaecker-Voirol et al., 2008). This formulation is compatible with the biodiesel norm EN14214. It can be incorporated without any problem in various blending biodiesel/ petrodiesel. The European norm EN590 valid in 2013 allows incorporating up to 7% FAME in diesel fuels. The construction of a demonstration plant for GTBE with a capacity of 20 ktpy was announced by the company GTBE N.V. in the Netherlands (www.procede.nl). A GTBE project proposed by CPS Biofuels Inc. in the United States received the Presidential Green Chemistry Challenge Award for 2012, the focus being on reducing up to 35% of particulates, improving the efficiency of E10 and E15 gasoline, as well as employed as icing inhibitor (FSII) in military and commercial kerosene-type jet fuel. GTBE are commercially available as CPS PowerShot since January 2011 (www2.epa.gov). A major problem in developing the pilot plant is the selection of raw materials for further scaling-up. From the above chemistry, it seems that good purity and water-free glycerol should be used. This is available directly from a solid catalyst process and with some effort from homogeneous catalysis processes. The second element is the source of the etherification agent. Isobutylene can be used in the context of a petrochemical complex. Tert-butyl alcohol (TBA) and isobutanol may be available from a biorefinery. In the last case, the GTBE would be 100% from renewable resources.

401

402

Chapter 11 FUEL ADDITIVES

11.2.1 Basis of Design 11.2.1.1 Chemical Reaction Network Fig. 11.1 shows the chemical reaction network for the esterification with isobutene. The reaction takes place in the presence of homogeneous (p-toluenesulfonic acid, pTSA) or heterogeneous (ion-exchange resins, zeolites) catalysts by successive addition of hydroxyl groups to the C]C double bonds, forming monoethers (two isomers), diethers (two isomers), and triethers. The reactions are equilibrium-limited. Although in view of incorporating a maximum hydrocarbon amount the target species would be the triether, in practice one gets a mixture of mono-, di-, and triethers. Because the monoethers are not fully soluble in diesel or biodiesel, the reaction is performed in a large excess of isobutene to shift the equilibrium toward the formation of diethers (around 65% w/w) and triethers. Moreover, the monoethers can be separated and recycled for further etherification. Glycerol is a trialcohol, with a high boiling point, and is very viscous. On the other side, isobutene is a hydrocarbon gas only slightly soluble in glycerol at higher pressure and temperature. Therefore, the reaction takes place in a two-phase system.

11.2.1.2 Physical Properties Behr and Obendorf (2003) showed that monoether and glycerol are completely miscible, while both diether and triether exhibit a miscibility gap with glycerol. Based on liquideliquid equilibrium experiments, Behr and Obendorf (2003) estimated the NRTL binary interaction of the NRTL model. Fig. 11.2 shows the ternary maps describing the liquideliquid equilibrium for a subsystem consisting of glycerol, monoether (ME), and diether (DE) and

H2C

OH

HC

OH

H2C

OH

+ H2C

C

H2C

O

HC

OH

H2C

OH

i-B

H2C

O

C4H9

HC

O

C4H9

H2C

OH

i-B

i-B

CH3

CH3 H2C CH H2C

G

C4H9

i-B

OH O OH

ME

C4H9

i-B

H2C

O

HC

OH

H2C

C4H9

H2C

O

C4H9

HC

O

C4H9

H2C

O

C4H9

i-B

O C4H9 DE

Figure 11.1 Etherification of glycerol with isobutene.

TE

Chapter 11 FUEL ADDITIVES

Figure 11.2 Ternary map for subsystem consisting of (A) monoether (ME), diether (DE), and glycerol; (B) isobutene (i-B), diether (DE), and glycerol.

for a subsystem consisting of glycerol, diether and isobutene (i-B). It can be observed that glycerol and diether are immiscible. When small amounts of monoether are added to the glycerolediether mixture, it is distributed between the two phases. One liquid phase exists when a large amount of monoether is added, while isobutene is miscible with both glycerol and diether. Fig. 11.3 analyzes the liquideliquid equilibrium at 25
C. Assuming that isobutene can be easily separated by a simple flash, the composition of a typical reactor outlet, point ME þ DE þ G falls in the single-phase region. This happens because the large amount of monoether increases the miscibility of glycerol and diether. The immiscibility can be exploited for separating the 1 G

xG

0.8

L1 L

0.6 0.4 0.2

L2 ME+DE+G

0 DE 0 0.2 0.4 0.6 0.8 x ME

1

ME

Figure 11.3 Liquideliquid immiscibility occurs when the reactor-outlet stream is mixed with fresh glycerol.

403

404

235.0

260.0 1.0 bar 0.5 bar 0.2 bar 1.0 bar 0.5 bar

210.0

Temperature / [°C]

T-x T-x T-x T-y T-y

185.0

235.0 210.0 185.0

Temperature / [°C]

260.0

Chapter 11 FUEL ADDITIVES

T-y 1.0 bar T-y 0.5 bar

T-y 0.2 bar

0.0

0.2

T-x 1.0 bar T-x 0.5 bar T-x 0.2 bar

T-y 0.2 bar

0.4

0.6

0.8

1.0

0.0

0.2

0.4

0.6

0.8

1.0

TE

DE

Figure 11.4 T-xy diagrams for the mono/diether and mono/triether mixtures.

reactants from products by mixing the fresh glycerol with the reactor outlet. The mixture (point L) separates into a glycerolrich phase (L1) and a DE-rich phase (L2). However, the DE-rich phase L2 contains significant amounts of ME. The volatility of isobutene is very high, compared to that of the glycerol and mono-, di-, and triethers. Therefore, it can be separated by a simple flash. T-xy diagrams for mono/diether and mono/triether mixtures (Fig. 11.4) show that the relative volatility is close to one at atmospheric pressure. However, it appears that at 0.2 bar separation by distillation is possible. In addition, vacuum distillation allows lower temperatures preventing products degradation. T-xy diagram (Fig. 11.5) for glycerol-diether shows a lowboiling azeotrope. However, the residue curve map (Fig. 11.5) of the DEeMEeG mixture shows only one distillation region where ME acts as a solvent for glycerol, allowing therefore high-purity diether to be obtained in one distillation unit.

11.2.1.3 Reaction Kinetics Behr and Obendorf (2003) carried out kinetic studies at 20 bar, 2% wt catalyst (p-toluenesulfonic acid), and 1000 rpm stirrer speed. The etherification reaction was performed at 343 K and 363 K. The ratio between glycerol and isobutene was 1:1 and 1:2. A simplified kinetic model was developed, which neglected the side reactions, the mass transfer processes, and phase equilibrium. The model uses overall concentration, meaning that component concentration in the reaction system was calculated from the concentration in the upper and lower phase. The Arrhenius equation

405

0 .0

5

Chapter 11 FUEL ADDITIVES

0 .1

5

0 .9

0. 2 5

250.0

0. 8

0. 3 5

0 .7

T-x 0.1 bar

0 .4

0. 6 5

0. 4

150.0

0.

55

E

0. 5

M

G

5

200.0

0. 6

0. 8 5

0. 2

100.0

0 .7

5

0 .3

Temperature / [C]

T-y 0.1 bar

0.

0 .9

5

1

0.0

0.25

0.5 G

0.75

1.0

0.05 0. 1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9 0.95

DE

Figure 11.5 T-xy diagram for the glycerolediether mixture; residue curve map of glycerol/mono/diether mixture (0.1 bar).

was used to describe temperature dependence of the reaction rates. Table 11.1 presents the kinetic parameters. k1

G þ i  B % ME; r1 ¼ k1 $CG $CiB  k 1 $CME k 1

(11.1)

k2

ME þ i  B % DE; r2 ¼ k2 $CME $CiB  k 2 $CDE k 2

(11.2)

k3

DE þ i  B % TE; r3 ¼ k3 $CDE $CiB  k3 $CTE k 3


 EA;j kj ¼ k0;j $exp  R$T

(11.3) (11.4)

Table 11.1 Glycerol Etherification With Isobutene. Kinetics Parameters. Preexponential Factors k0;1 k0;1 k0;2 k0;2 k0;3 k0;3

1

1

L$min $mol min1 L$min1$mol1 min1 L$min1$mol1 min1

Activation Energies/(kJ/kmol)

3.04  10 3.69  1013 1.70  1011 8.54  1014 2.26  1010 6.35  1015 8

EA;1 EA;1 EA;2 EA;2 EA;3 EA;3

74.04 111.78 92.80 118.06 92.56 125.13

406

Chapter 11 FUEL ADDITIVES

11.2.2 Conceptual Design Several alternative processes for the manufacturing of GTBE by etherification with isobutene are described in the literature (Fig. 11.6). The reaction may take place, at high pressure and using pTSA as catalyst, in a CSTR (Gupta, 1995; Cheng et al., 2011; Vlad et al., 2013) or a series of CSTRs (Behr and Obendorf, 2003). Alternatively, a solid catalyst can be used (Noureddini, 2000; Di Serio et al., 2010). Isobutene is separated from the reaction mixture as vapor in a flash or a striping column. Unreacted glycerol and monoether are separated of di- and triethers mixture in two different ways: (1) by distillation (Behr and Obendorf, 2003; Cheng et al., 2011; Vlad et al., 2013) and (2) by extraction with water (Gupta, 1995; Noureddini, 2000), glycerol (Behr and Obendorf, 2003; Cheng et al., 2011; Vlad et al., 2013), or biodiesel (Di Serio et al., 2010) and recycled. Final purification of high ethers is achieved by vacuum distillation.

11.2.3 Process Design In the following, details of a glycerol etherification plant (Vlad et al., 2013) will be presented. Pure glycerol (G0) and a mixture of 90% isobutene and 10% isobutane (IB) were considered as raw materials. The flowsheet is shown in Fig. 11.7, together with the main sizing elements. The etherification of glycerol with isobutene takes place in a CSTR. The reaction temperature and pressure are set to 90
C and 14 bar, respectively, when the reaction mixture is liquid. Because of the large amounts of ME, the reaction mixture comprises of only one liquid phase. To exploit the immiscibility of the less polar components (C4 hydrocarbons, di- and triethers) in glycerol and monoether, the first separation step is performed by an extraction column (EX). The raffinate stream 2a contains isobutene, inert, DE, TE, and small quantities of G and ME and is sent to separation. Column C1 separates the isobutene and the isobutane. The column has a total condenser and is operated at 5 bar to keep the isobutene as liquid. The entire amount of isobutene and isobutane present in the feed is recovered in stream “3c.” A purge is necessary to remove the inert from the plant. The recycle (stream 3d) is mixed with fresh isobutene and fed to the reactor. Column C2 separates di- and triether from glycerol and monoether. The column has a total condenser and is operated under vacuum (0.2 bar) to avoid high temperature in the bottom of the column. The distillate “4a” is the plant product. The bottom stream, containing mainly G and ME, is mixed with

Chapter 11 FUEL ADDITIVES

407

Figure 11.6 Processes for etherification of glycerol with isobutene.

408

Chapter 11 FUEL ADDITIVES

4b Purge (C4)

3d Recycle (C4)

IB0 MIXER

Qc = 17.08 kW

1b 1

5 bar

C4

3c

RR = 0.18 90 14 bar REACTOR 16.1 m3

14 bar

5 bar

C1

Qc = 207.02 kW NT = 10 D = 0.12 m

2a 1a

G0

0.2 bar

QR = 100.95 kW

2 EX 2 stages

Glycerol and mono-ether

Glycerol and ethers

3a

C2

RR = 3.4

4a Di- and tri-ether

NT = 17 D = 0.47 m

QR = 136.3 kW 3b

Figure 11.7 Flowsheet of the glyceroleisobutene etherification plant.

the glycerol-rich phase from the extraction column and fed to the reactor. Table 11.2 presents the mass balance of the plant. The economic calculations show total annual cost of 321$103 US$/year (payback period 3 years). The investment and operating costs are 670$103 US$ and 74$103 US$/year, respectively. It should be noted that several other designs are possible, depending on the reactor volume and the ratio of isobutene/ glycerol at reactor inlet. The reactor volume can be reduced, but this has a negative impact on the economics because the costs of separation increase. Lower costs can be achieved by operating the plant at lower ratio of the reactor-inlet streams (C4 to glycerol and monoether) because isobutene recycling costs are reduced. However, the controllability of the plant becomes worse, as the sensitivity to various disturbances (such as throughput, purity of the fresh isobutene, catalyst activity, reaction temperature, performance of the extraction unit) is higher (Vlad et al., 2013).

11.3 Glycerol Etherification With Tert-butanol The etherification can employ also alcohols, as tert-butanol (TBA). The reaction network is similar to the one shown in Fig. 11.1, except that one molecule of water is obtained in each

Chapter 11 FUEL ADDITIVES

409

Table 11.2 Stream Results for the Process of Glycerol Etherification With Impure Isobutene (Vlad et al., 2013). Stream Name

G0

IB0

1a

1b

2

2a

3a

3b

3c

4a

4b

Flow/(kmol/hr) Flow/(kg/hr) Temperature/(o C) Pressure/(bar) Flow rate/(kmol/hr) Glycerol Isobutane Butene Monoether Diether Triether

2.25 207.5 90 14

6.15 346.8 20 14

5.22 598.8 104.5 14

7.78 439.6 24.2 14

8.08 1038.4 90 14

5.52 697.4 73.1 5.5

4.81 548.5 102.8 14

0.40 50.3 206.9 0.2

2.85 163.0 39.6 5

2.25 481.7 158 0.2

1.23 70.3 39.6 5

2.25 0 0 0 0 0

0 5.5 0.61 0 0 0

2.99 0.12 0.12 1.74 0.19 0.04

Trace 6.35 1.43 Trace Trace 0

0.74 1.55 1.55 1.74 2.02 0.46

0.19 1.43 1.43 0.21 1.83 0.42

2.80 0.12 0.12 1.53 0.18 0.04

0.19 Trace Trace 0.20 0.009 Trace

Trace 1.42 1.42 Trace Trace Trace

0.005 0.004 Trace 0.002 1.82 0.42

Trace 0.61 0.61 Trace Trace Trace

etherification step. In this case, the compatibility of reactants is much better because they belong to the same chemical class. However, the cost of TBA is obviously higher than that of isobutene. Moreover, water and TBA form a low-boiling, homogeneous azeotrope (0.621 mol TBA, 79.97
C; 1 bar, NRTL model with parameters estimated by UNIFAC). This can be broken by using a suitable solvent, but the separation section is more complex. Therefore, despite the advantage of a lower operating pressure, etherification with TBA in a conventional reactoreseparatione recycle (RSR) process is penalized by the cost of separations. However, this reaction can be easily performed by reactive distillation (RD) (Vlad and Bildea, 2012), as it will be shown later.

11.3.1 Basis of Design 11.3.1.1 Reaction Stoichiometry and Kinetics The etherification of glycerol with TBA proceeds according to the following reactions: G þ TBA%ME þ H2 O ME þ TBA%DE þ H2 O DE þ TBA%TE þ H2 O

410

Chapter 11 FUEL ADDITIVES

Table 11.3 Glycerol Etherification With Tert-butyl Alcohol (TBA). Kinetic Parameters. Reaction

Equilibrium Constant

Ka , 70
C

Rate Constant (mol sL1 kgL1)

G þ TBA%ME þ H2 O ME þ TBA%DE þ H2 O DE þ TBA%TE þ H2 O

Keq1 ¼ expð2:581  754:8=TÞ Keq2 ¼ expð1:228  942:1=TÞ Keq3 ¼ expð1:779  2212=TÞ

1.46 0.22 0.009

k1 ¼ expð17342  6835=TÞ k2 ¼ expð26953  10382=TÞ k3 ¼ expð26953  10382=TÞ

DE, diether; G, glycerol; ME, monoether; TE, triether.

Small amounts of isobutene are formed because of TBA dehydration: TBA%IB þ H2 O
cova
et al., Suitable catalysts are ion-exchange resins (Klepa 2005, 2006; Yusof et al., 2008), p-TSA and sulfuric acid (Yusof et al., 2008), and silica-supported acids (Frusteri et al., 2009). Reaction temperatures in the range 303e363
C lead to glycerol conversions exceeding 64% and mono- to diether ratio in the range 4:1 to 6:1 (Chang and Chen, 2011). Amberlyst-15 showed the highest activity. Kiatkittipong et al. (2011) derived a kinetic model for glycerol etherification with TBA, with parameters obtained by regressing measured data from an autoclave reactor (Table 11.3).
 1 r1 ¼ k1 xG xTBA  x ME x W (11.5) Keq;1
 1 x DE xW (11.6) r2 ¼ k2 x ME x TBA  Keq;2
 1 x TE x W r3 ¼ k3 x DE x TBA  (11.7) Keq;3 Yusof et al. (2008) report glycerol conversions exceeding 64% and mono- to diether ratio in the range 4:1e6:1, obtained using Amberlyst-15, Amberlite IR-120, Montmorillonite K10, p-toluenesulfonic acid, and sulfuric acid as catalysts. Frusteri et al. (2009) studied the etherification of glycerol with TBA in presence of lab-made silica-supported acid catalysts. Experiments were carried out in batch mode at temperature ranging from 303 to 363 K. Ozbay et al. (2010) compared various solid acid catalysts, such as Amberlyst-15, Amberlyst-16, and Amberlyst-35,

Chapter 11 FUEL ADDITIVES

Nafion-SAC-13, and gamma-alumina. Amberlyst-15 showed the highest activity at about 110 degrees C, while A-16 gave higher diether selectivity values. Chang and Chen (2011) present a systematic optimization of the glycerol etherification with TBA based on the small-scale experimental data. Kiatkittipong et al. (2011) performed experiments in an RD column and compared with Aspen Plus simulation results. It should be remarked that the column worked more as a series of CSTRs (the trays) with the reactants fed in countercurrent because the distillate stream contained a mixture of water and TBA, while the bottom stream contained TBA, glycerol, and mono-, di-, and triether. Therefore, separation of reactants and products from the column-outlet streams was still necessary. Ozbay et al., 2010 investigated the liquid-phase etherification of glycerol with TBA in a continuous flow reactor using Amberlyst-15 as catalyst. In a different reaction pathway, Al-Lal et al. (2012) suggest transformation to epichlorohydrin followed by etherification with TBA.

11.3.1.2 Physical Properties Table 11.4 presents the NRTL binary interaction parameters between the components involved in glyceroleTBA etherification process, obtained from three sources: binary parameters between glycerol and ethers were taken from Behr and Obendorf (2003) experimental study, binary parameters between TBA and ethers were estimated by UNIFAC-LL, and binary parameters between TBA, water, and glycerol were taken from Aspen Plus database (based also on experiments).

11.3.1.3 Vapor Liquid Equilibrium Table 11.5 presents the boiling points of the main components and their azeotropes. Small amounts of i-B that are formed by TBA dehydration can be easily removed because of lower boiling point. The separation of TBA and water from glyceroleethers mixtures appears to be easy. Also, glycerol and monoether, which are recycled, can be obtained as the bottom product of a distillation column operating under vacuum. Obtaining high-purity DE product seems difficult because of the low-boiling GeDE azeotrope. However, the residue curve map of the DEeMEeG mixture shows only one distillation region where ME acts as a solvent for glycerol, allowing therefore high-purity diether to be obtained in one distillation unit. TBA and water form a low-boiling homogeneous azeotrope. This can be broken by using a suitable solvent, for example, 1,4-butanediol.

411

412

Chapter 11 FUEL ADDITIVES

Table 11.4 Glycerol Etherification With Tert-butyl Alcohol (TBA). NRTL Binary Parameters. Component

Parameter/(K)

i

J

bij

bji

Source

G G G G G G TBA TBA TBA TBA TBA ME ME DE

TBA ME DE TE H20 i-B ME DE TE H20 i-B DE TE TE

6.51 207.34 1573.30 1573.30 170.91 721.75 185.75 72.77 5.21 1372.38 630.84 630.83 630.83 93.94267

408.59 79.22 528.53 528.53 272.6 937.02 220.48 67.17 426.97 303.41 48.13 680.40 680.40 89.09376

UNIFAC-LL Behr and Obendorf Behr and Obendorf Behr and Obendorf Behr and Obendorf Behr and Obendorf UNIFAC-LL UNIFAC-LL UNIFAC-LL Aspen Plus database UNIFAC-LL Behr and Obendorf Behr and Obendorf UNIFAC-LL

DE, diether; G, glycerol; i-B, isobutene; ME, monoether; TE, triether.

Table 11.5 Glycerol Etherification With Tert-butyl Alcohol (TBA). Boiling Point for Pure Components and Azeotropes (P [ 1 bar). Component//Azeotrope

T/(
C)

Destination

Isobutane TBA (0.6209)/water (0.3791) TBA Water G (0.1951)/DE (0.8049) DE ME G

6.25 79.97 82.42 100 233.5 240.4 256.61 287.85

By-product To TBA/water separation Recycle By-product e Product Recycle Recycle

DE, diether; G, glycerol; ME, monoether.

0.4

0.5

0.3

0.7

0.4

0.6

0.5

R TE

WA

nt

0.6

So lve

0.7

0.3

0.8

0.2

0.9

0.1

Chapter 11 FUEL ADDITIVES

0.8

0.2

C3

0.1

0.9

EX C4 0.1

0.2

0.3

0.4

0.5 TBA

0.6

0.7

0.8

0.9

Figure 11.8 Residue curve map of the tert-butyl alcohol (TBA)ewatere1,4-butanediol mixture.

Fig. 11.8 shows the residue curve map of the TBAewatere1,4butanediol mixture, indicating also the possibility of breaking the watereTBA azeotrope. Thus, the initial watereTBA mixture is firstly separated into TBA and azeotrope (column C3). The azeotrope enters in the lower part of the extractive distillation column (EX), while the solvent is fed at the top. The distillate contains water, while the bottom stream consists of solvent and TBA, which is easily separated in column C4.

11.3.2 ReactoreSeparationeRecycle Process 11.3.2.1 Plant Flowsheet Fig. 11.9 presents the flowsheet of the reaction section, while Fig. 11.10 details the azeotrope separation section. The control loops are also depicted. Detailed stream reports are presented in Table 11.6 and in Table 11.7. The etherification of glycerol with TBA takes place in a plug flow reactor in presence of 400 kg Amberlyst. The reaction temperature and pressure are set to 70
C and 5 bar, respectively, where the reaction mixture is liquid. The reactor-outlet stream is

413

TBA0

Azeotrope separation

3a

LC

FC TBA1

r1

W

PC

x 1

FC

LC G0

PFR FC PC

GM1

C1 LC 2 TC LC D+T 3b

C2

TC

XC

TC LC

Figure 11.9 Glycerol etherification with tert-butyl alcohol (TBA). Flowsheet and control loops of the reaction section.

W XC

IB

PC

S0 LC PC PC TC LC

EX LC TC AZ LC TBA1b

TC

C3

3a

TC XC

TBA+S

C4

TC LC TC

LC S

TBA1a

Figure 11.10 Glycerol etherification with tert-butyl alcohol (TBA). Flowsheet and control loops of the azeotrope separation section.

Chapter 11 FUEL ADDITIVES

415

Table 11.6 Glycerol Etherification With Tert-butyl Alcohol (TBA). Stream Results for the Reaction Section (Fig. 11.9). Stream Name

G0

TBA1

TBA0

GM1

1

2

3a

3b

W

DDT

Flow/(kmol/hr) Flow/(kg/hr) Temperature/(o c) Pressure/(bar) Flow/(kmol/hr) G TBA ME DE TE Water Solketal Isobutane

2.15 198 20 1

26.5 1883 79.8 5

4.37 324 25 1

6.23 759 140.9 5

32.75 2642 70 5

32.8 2642 70 4.95

26.47 1635 29.3 1.2

6.32 1006.6 177 0.15

4.35 80.8 25 1

2.24 445.6 139.5 0.1

2.15 0 0 0 0 0 0 0

Trace 24.9 Trace Trace Trace 1.5 0.1 0

0 4.37 0 0 0 0 0 0

2.94 Trace 3.27 0.01 Trace Trace 0 0

2.94 24.9 3.27 0.01 Trace 1.5 0.1 Trace

0.8 20.6 3.27 2.14 0.002 5.81 0.1 0.04

Trace 20.6 Trace Trace Trace 5.81 Trace 0.04

0.81 0.002 3.27 2.14 0.002 0.005 0.1 Trace

0 0.04 0 0 0 4.31 Trace 0

0.01 0.003 0.007 2.13 0.0019 0.005 0.1 Trace

DE, diether; G, glycerol; ME, monoether; TE, triether.

Table 11.7 Glycerol Etherification With Tert-butyl Alcohol (TBA). Stream Results for the Azeotrope Separation Section (Fig. 11.10). Stream Name

3a

TBA1a

Az

S0

W

TBA D S

IB

TBA1b

S

Flow/(kmol/hr) Flow/(kg/hr) Temperature/(o C) Pressure/(bar) Flow/(kmol/hr) TBA Isobutane Solketal Water

26.46 1635 29.3 1.2

11.7 858.5 97.8 1.8

14.7 776.6 76 1.0

90.6 8164.5 30 1.2

4.35 80.8 25 1

100.9 8858 156 1.5

0.04 2.22 25 1

10.38 700.1 80 1.0

90.5 8158 231 1.13

20.61 0.04 0 5.8

11.5 Trace 0.0 0.2

9.1 0.04 Trace 5.57

0 0 90.6 0

0.04 0.003 Trace 4.31

9.05 Trace 90.6 1.2

Trace 0.039 0 0.001

9.03 Trace 0.1 1.2

Trace Trace 90.5 Trace

416

Chapter 11 FUEL ADDITIVES

routed to column C1. TBA and water are separated as top product, while a mixture of glycerol and ethers leaves the column as bottom product. The column is operated under vacuum (0.1 bar) to avoid high temperature in the bottom of the column. Column C2 separates the mixture of di- and triethers. The bottom product contains glycerol and monoether and is recycled. The column is operated under vacuum (0.1 bar) to avoid product degradation. Column C3 separates the TBA/water azeotrope from TBA, which is mixed with fresh TBA and recycled. The extractive distillation column (EX) is fed on the bottom with TBA/water azeotrope and at the top with 1,4-butanediol, which is the solvent. The solvent extracts TBA and is eliminated on the bottom of the column, while the water is removed as liquid at the top. The column has a partial condenser to eliminate isobutene traces. Column C4 recovers the solvent. TBA is removed at the top of the column, is mixed with TBA stream from column C1 and with fresh TBA, and is recycled. Table 11.8 presents the economic evaluation of the design, for a payback period of 3 years.

11.3.2.2 Dynamics and Control The control structure investigated here makes it typical for processes with two reactants, which are distinctly recycled (Dimian et al., 2014). The flow rate of fresh glycerol is set to the value FG,0, which also sets the amount of product obtained F4 (F4 ¼ FG,0). Moreover, the ratio of reactants at reactor inlet, r1 ¼ FTBA,1/FGþME,

Table 11.8 Glycerol Etherification With Tert-butyl Alcohol (TBA). Economic Evaluation Results. Reactor Diameter (m) 0.35 Reflux ratio Diameter (m) No. of trays Height (m) 2 Reboiler duty (kW) Condenser duty (kW) Cost ($) 9296 Cost ($)

Column C1 Column C2 Column C3 Column EX Column C4 0.32 0.85 6 343 348.7 297,966

Energy cost ¼ 291,663 $/year Equipment cost ¼ 2,440,746 $ Total annual cost ¼ 1.10  103 $/year (3 years payback)

1.83 0.55 15 170 168.6 221,775

4 0.7 41 716.2 654 589,424

1.5 0.5 30 525 170 759,704

2 0.55 7 563.3 324.4 338,157

Chapter 11 FUEL ADDITIVES

TBA

TBA recycle

3b

0

417

1

LC

r1 = 4

0.8 400

1b

Glyc erol 0

3a

XG

FC

x

to Reactor

m cat / [kg] =200

0.2 0

LC

0

FC

1

3

0.4

m

4 0.8

5

X TBA

r1= 2

0.4

cat

/ [kg] = 400

r1 = 2

0.3

3 0.6

4

F G,0 / [kimol/h]

1

XG

2

to Reactor

1a

300

0.4

Glycerol and ethers recycle

FC

0.6

3 4

0.2 0.1

0.2 m cat / [kg] = 400 0

0 0

1

2

3

4

5

0

1

F G,0 / [kmol/h]

2

3

4

5

F G,0 / [kmol/h]

Figure 11.11 Glycerol etherification with TBA. Reactants conversions versus fresh glycerol flow rate, for different values of catalyst mass and ratio between reactor-inlet flow rates.

is also fixed. Fig. 11.11 presents the strategy of feeding the reactants and the results of a sensitivity analysis which considered a kinetic reactor and perfect separation (Vlad et al., 2012). Glycerol conversion versus fresh glycerol flow rate is plotted for different amounts of catalyst. Different amounts of glycerol can be processed, up to maximum value which increases with the catalyst mass. It can be seen that 400 kg of catalyst allows doubling the production rate. For 400 kg of catalyst, the lower diagrams show that the glycerol and TBA conversions have small sensitivity to the production rate or the ratio between reactor-inlet flow rates. Fig. 11.12 presents results of dynamic simulation, performed in Aspen Plus Dynamics. Molar and mass flow rates together with mass fractions are shown. Starting from the steady state, two disturbances were introduced. At time 2 h, a 10% wt. water impurity in the fresh glycerol was introduced. Later, the flow rate of fresh glycerol was increased to 2.15 kmol/h, at time 40 h. The results demonstrate that the nominal operating point is stable, and the plant achieves stable operation when disturbances are introduced.

418

Chapter 11 FUEL ADDITIVES

Figure 11.12 Glycerol etherification with tert-butyl alcohol (TBA). Dynamic simulation results.

11.3.3 Reactive Distillation for Glycerol Etherification with Tert-butyl Alcohol In the following, a glyceroleTBA etherification process, employing RD, is presented (Vlad and Bildea, 2012). This alternative produces high-purity glycerol ethers and water at a greatly reduced cost. The controllability of the process is proven by dynamic simulation.

11.3.3.1 Reactive Distillation Design Fig. 11.13 shows the RD column performing the etherification of glycerol with TBA. The column has a diameter of 0.3 m and 30 trays. There are 10 reactive stages, each holding 40 kg of catalyst. The column is fed with 2.25 kmol/h of glycerol heated at 150
C and with 4.51 kmol/h of vapor TBA. The reboiler duty is 93.4 kW, while 86 kW is necessary for reactant preheating. The top product contains water and small amounts of GTBE and TBA. A fraction of the condensate is recycled (reflux ratio ¼ 0.5). The rest is sent to a liquideliquid separator, where the organic phase is recycled, and the aqueous phase is removed as product. The bottom

Chapter 11 FUEL ADDITIVES

419

Qc = 78.9 kW

xW = 97.5 %wt. xDTBG = 2.26 %wt

RR = 0.5 1

2.25 kmol/h 150 G

0.1 kmol/h xW = 47.3 %wt. xDTBG = 52.0 %wt.

10

4.511 kmol/h xW = 99.8 %wt. Water

Recycle 4.511 kmol/h 95

diam = 0.3 m mcat = 400 kg

20

TBA

30

2.25 kmol/h xDTBG = 99.2 %wt. xTTBG = 99.2 %wt. xG = 0.03 %wt.

GTBE

QR = 93.4 kW

Figure 11.13 Flowsheet of etherification of glycerol with tert-butanol alcohol by reactive distillation.

product contains DE with small amounts of TE and traces of glycerol. Fig. 11.14 presents temperature, reaction rate, and the liquid concentration profiles along the RD column, obtained from simulation in Aspen Plus. The reaction takes place in the reactive part of the column (trays 10 to 20), where GTBE are formed. The maximum concentration of GTBE and water are observed on stage 30 and 1, respectively. 1.0

0.8

250

0.7

0.4 0.3

T

100

0.2

GTBE Glycerol

0.6

x

o

T / [ C]

0.5

r

150

r / [kmol/h]

0.6

200

Water

0.8

0.4

TBA

0.2

0.1 50

0 0

5

10

15

20

Stage number

25

30

0.0 0

5

10

15

20

25

30

Stage number

Figure 11.14 Temperature, reaction rate (left), and liquid mass fraction (right) profiles along reactive distillation column.

420

Chapter 11 FUEL ADDITIVES

1

r = 2.005

0.999

2.000

0.998

1.995

40 0.998

x GTBE,B

x GTBE,B

1

0.997

60

30

0.996 m / [kg] = 20

0.994 0.992

0.996

0.99

0.995 0.3

0.4

0.5

0.6

0.7

0.8

0.4

RR

0.5

0.6

0.7

0.8

RR

Figure 11.15 Mass fraction of glycerol tert-butyl ethers (GTBE) in product versus the reflux ratio, for different tert-butyl alcohol (TBA): glycerol ratio (left) and amount of catalyst on each reactive stage (right).

Fig. 11.15 (left) presents the influence of reflux ratio (RR) on the product purity for different values of the TBA/glycerol feed rate. Although higher purity can be achieved at reflux ratios around RR ¼ 0.35, multiple steady states exist in this region, which could lead to operating difficulties. Therefore, the reflux ratio was set to RR ¼ 0.5, for which a single steady state exists. A small excess of TBA is necessary for high purity of the GTBE product. The right diagram shows the influence of the amount of catalyst on each tray. As expected, higher purity is obtained when more catalyst is used. Above 40 kg/tray, the amount of catalyst has a small influence on product purity.

11.3.3.2 Dynamics and Control Fig. 11.16 shows the control of the RD column. The pressure and levels are controlled by conventional loops. The reflux ratio is kept constant at the design value. The reboiler duty is ratioed to the glycerol feed rate. Sensitivity analysis (Fig. 11.15) shows that feeding the reactants in the stoichiometric ratio is necessary to ensure high product purity. However, a simple feed-forward scheme would not work because of unavoidable measurement and implementation errors. Therefore, the concentration controller XC provides the necessary feedback. Fig. 11.17 presents the results of dynamic simulation. The simulation starts from the nominal steady state, which is maintained for the first hour. Then, a disturbance in column feed was introduced: at time t ¼ 1 h, the glycerol feed rate was increased from 2.25 kmol/h to 2.5 kmol/h (left). The TBA flow rate is adjusted to the required stoichiometric ratio. Consequently, the GTBE flow rate increases, while the products purity remains constant. A new steady state is established in about 1 h. The right diagrams show the effect of a 10% error in glycerol

Chapter 11 FUEL ADDITIVES

421

PC

LC1

X x

GLY

Decant

1

RDCOL

LC2

LC3

RR RECYCLE

WATER

TBA

r

X

x

LC4 GTBE Sum Σ

Σ

X

x

QR

XC

Figure 11.16 Control structure of glycerol etherification process by reactive distillation.

Figure 11.17 Dynamic simulation results for glycerol etherification plant with TBA by RD. Feed and product flow rates and product purity (mass fraction) for 10% increase of the production rate (left) and glycerol flow rate inaccuracy (right).

422

Chapter 11 FUEL ADDITIVES

Table 11.9 Economic Evaluation Results for Glycerol Etherification Process With Tert-butanol Alcohol by Reactive Distillation (RD). Main Equipment RD Column Condenser Duty ¼ 79 kW No. of trays ¼ 28 Cost ¼ 46,543 $ D ¼ 0.3 m H ¼ 20.1 m Cost ¼ 61,337 $ Total annual cost ¼ 87,387 $/year Energy cost ¼ 28,383 $/year

Reboiler Duty ¼ 93.6 kW Cost ¼ 61,190

Decanter D ¼ 0.4 m H ¼ 0.8 m Cost ¼ 4941 $

flow rate measurement occurring at time t ¼ 1 h and t ¼ 10 h, respectively. Small fluctuations on GTBE and TBA flow rates and on products purity are observed. However, the purity of GTBE remains high, and the degradation of water purity is very small.

11.3.3.3 Economic Evaluation Table 11.9 presents the size and capital cost for the main equipment of the plant, together with operating costs. Compared to the conventional RSR process, using the RD column leads to reduction of 95% in the capital cost and 97% in the energy cost.

11.4 Glycerol Ketalization Solketal (2,2-dimethyl-1,3-dioxolane-4-methanol) can be used as a fuel additive to reduce the particulate emission and to improve the cold flow properties of liquid transportation fuels (Silva et al., 2010). Solketal is obtained by catalytic condensation of glycerol with acetone, in the presence of an acid catalyst (Fig. 11.18). Despite several research efforts, the study of this potentially important reaction was limited to laboratory scale. In this following, the design of a solketal plant is presented. Next section presents the reaction equilibrium and kinetics and the physical properties of the species involved in the process. Consideration of the Reaction-Separation-Recycle (RSR) structure of the plant allows performing a preliminary mass balance. Then, details of

Chapter 11 FUEL ADDITIVES

H3C

+ H3C

CH3

OH

O

OH

HO

O

O

+ H2O

CH3 OH

Acetone (A)

Glycerol (G)

Solketal (S)

Water (W)

Figure 11.18 Glycerol ketalization.

the reactor and separation units, together with detailed stream report, equipment sizing, and operating conditions are given. A plantwide control structure is suggested and evaluated by rigorous dynamic simulation. Finally, the economic evaluation provides the total annual cost.

11.4.1 Basis of Design In the following, we will present the design and evaluation of a solketal plant processing 296 kg/h of glycerol, assumed to be the by-product of a 37,200 tones/year biodiesel plant (Zaharia et al., 2015). The initial data include information about the equilibrium and kinetics of the chemical reaction and the physical properties of the species involved in the process. The reaction follows an LHHW mechanism where two active centers are involved. When Amberlyst-35 is used as catalyst, the following relationships describe the reaction kinetics:
 1 cS cW cG cA 1  K c cG c A r ¼ k (11.8) 2 ð1 þ KW cW Þ where subscripts G, A, S, and W refer to glycerol, acetone, solketal, and water, respectively. Experiments performed in a batch reactor, using 1% wt catalyst/glycerol, gave the following kinetic parameters (Nanda et al., 2014):

  55600 kJ=kmol 1 1 3  m3 kmol=s k ¼ 3:98  10 exp  R T 308 (11.9) 3615:4 T 7785:8 3  m kmol ¼ 25:1925 þ T

ln Kc ¼ 11:308 þ ln Kw

(11.10) (11.11)

423

424

Chapter 11 FUEL ADDITIVES

Table 11.10 Boiling Points of Chemical Species and Their Azeotropes (1 atm). Component

Boiling Point

Acetone Binary azeotrope, homogeneous Ethanol (0.9013); water (0.0987) Ethanol Water Solketal Glycerol

56.14
C (saddle) 78.18
C (saddle) 78.31 (saddle) 100
C (saddle) 189
C (saddle) 287.7
C (stable node)

To improve the solubility of glycerol in acetone and the homogeneity of the solution, ethanol is added to the reaction mixture (Nanda et al., 2014). The boiling points of chemical species involved in the process and their azeotropes are presented in Table 11.10.

11.4.2 Conceptual Design The RSR structure of the plant is presented in Fig. 11.19. Acetone (A) and ethanol (E) are the lightest components of the reactoroutlet mixture; therefore, they are recycled together. We note the difficult separation of high-purity acetone from acetone-water mixture and the formation of the ethanolewater azeotrope; therefore, this recycle stream contains significant amounts of water (W). Solketal (S) and water are the intermediate boiling components which can be easily separated from glycerol (G). The large difference between the boiling points allows obtaining highpurity products and glycerol recycle. To obtain a preliminary mass balance of the plant, the conversion of the equilibrium-limited reaction must be specified. Fig. 11.20 presents the temperature dependence of the equilibrium conversion, calculated for different reactor-inlet acetone/ glycerol (MAG) and ethanol/glycerol (MEG) molar ratios. The calculations took into account the amount of water in the reactor feed because of ethanolewater azeotrope. Because the reaction is exothermal, the conversion decreases with temperature. Glycerol conversion is favored by large acetone excess. Adding ethanol as ethanolewater mixture has an opposite effect because water is also a reaction product.

425

Chapter 11 FUEL ADDITIVES

A: 21.8 kmol/h; E: 4.2 kmol/h; W: 0.9 kmol/h

A: 3.2 kmol/h E: G: 3.2 kmol/h

A: 25 kmol/h E: 4.2 kmol/h W: 0.9 kmol/h G: 4.2 kmol/h

Reactor XG = 0.765

A: 21.8 kmol/h G: 1.0 kmol/h E: 4.2 kmol/h W: 4.1 kmol/h S: 3.2 kmol/h

S: 3.2 kmol/h Separation

W: 3.2 kmol/h

G: 1 kmol/h

Figure 11.19 Preliminary mass balance of glycerol ketalization plant (A e acetone; G e glycerol; E e ethanol; W e water; S e solketal). 1

0.9

XG,eq

XG,eq 0.8

30

40

50

MEG = 1 0.8

MAG = 4

20

MAG = 6

0.9

MAG = 8 MAG = 6

0.7

1

MEG = 1

60

T / [°C]

0.7

MEG = 5 MEG = 10 20

30

40

50

60

T / [°C]

Figure 11.20 Glycerol equilibrium conversion versus reaction temperature, for different values of the acetone/ glycerol ratio MAG (left) and ethanol/glycerol ratio MEG (right).

Based on results presented in Fig. 11.20, we choose the following reaction conditions: 35
C; acetone/glycerol ratio 6; ethanol/glycerol ratio 1. The design aims for glycerol conversion XG ¼ 0.765, which is about 0.85% of the equilibrium value. These data allow performing the preliminary balance of the plant. Considering the separation section as a black box which provides the product streams (solketal, water) and the recycles (glycerol and acetoneeethanolewater), the component flow rates in feed, recycle, and product streams can be calculated. A summary of the results is presented in Fig. 11.19.

11.4.3 Design of the Chemical Reactor Because of the small amount of catalyst that is necessary (1% wt. catalyst/glycerol), the reaction will be performed in a stirred

426

Chapter 11 FUEL ADDITIVES

tank reactor, where mixing improves the mass transfer of the reactants from the liquid bulk to the catalyst surface. To account for external and internal resistance to mass transfer, we amended the kinetic constant by an effectiveness factor h ¼ 0.2. The catalyst can be retained inside the reactor by placing it in a basket, separation in a cyclone, or an external decanting unit. The mathematical model consists of mass balance Eqs. (11.12) and (11.13). The species concentrations ck from the reaction rate (11.8) were calculated by assuming the following molar volumes (in m3/kmol): 0.0732 (glycerol), 0.0740 (acetone), 0.0582 (ethanol), 0.0181 (water), 0.1243 (solketal). The reactor-inlet flow rates Fk,1 were taken from the preliminary mass balance (Fig. 11.19). A conversion XG ¼ 0.765 was specified. Solution of Eqs. (11.12)e(11.15) leads to a reaction volume V ¼ 3.4 m3. To account for performance degradation because of additional amount of water being recycled, a volume of 4 m3 will be used. Fk;2  ðFk;1 þ nk $xÞ ¼ 0; k ¼ G; A; E; W; S x  V$h$rðck Þ ¼ 0 XG  ck ¼ P k

FG;1  FG;2 ¼ 0 FG;1

Fk;2 ; k ¼ G; A; E; W; S V mk $Fk;2

(11.12) (11.13) (11.14) (11.15)

11.4.4 Design of the Separation Section To design the separation section, the species found in the reactor effluent were ordered by boiling points (Table 11.10) and the structure of the separation section was decided (Fig. 11.21). The separation-outlet streams are the product (solketal), by-product (water), and two recycles (acetoneeethanolewater and glycerol). Distillation was chosen as separation method, following the “lights out first” heuristic. Each column was designed using the rigorous RADFRAC model from Aspen Plus. The first column (COL-1) separates the light components (acetone and ethanol) from the heavy ones (water, solketal, and glycerol). Besides acetone (0.811 M) and ethanol (0.156), the distillate contains some amounts of water because of azeotrope formation. The bottom stream, containing water, solketal, and glycerol, is sent to the second column (COL-2), from which water is obtained as distillate (0.999 purity). The remaining mixture is further

Chapter 11 FUEL ADDITIVES

Separation section

Reaction section MIX-A

Acetone

A-REC

A-0

Ethanol

B10

E-0

REACTOR

RIN

Water

HEATER

Glycerol

427

W ROUT

MIX-G

S

G-0 COL-2

COL-1 WSG

COL-3 SG

G-REC

Figure 11.21 Flowsheet of the glycerol ketalization plant.

separated in column COL-3 to provide the solketal product as highpurity distillate and the glycerol recycle as bottoms. The tray-sizing facility of Aspen Plus was used to find the columns diameter. The height of the columns was calculated considering 0.6 m tray spacing and allowing 20% for the top and bottom parts. Aspen Plus simulation also provided the reboiler and condenser duties. The condenser and reboiler areas were estimated considering a heat transfer coefficient of 500 W/m2/K and a temperature difference of 20 degrees. The flow rate of cooling water was calculated considering a 10 degrees temperature increase. The details of the separation section and the distillation columns sizing are presented in Table 11.11. Detailed stream results are given in Table 11.12.

11.4.5 Plantwide Control The plantwide control scheme is shown in Fig. 11.22. The glycerol feed rate is given by an upstream unit. The ratio between the reactor-inlet acetone and glycerol flow rates is set in a feed-forward manner. The amount of fresh acetone is set by the level controller of the MIXER vessel. Reaction temperature and reactor levels are controlled by the duty and outlet flow, respectively. Control of the distillation columns is standard: condenser duty controls the pressure; levels in the reflux drum and column sump are controlled by the distillate and bottoms rate, respectively; a temperature in the stripping zone is controlled by the reboiler duty; the reflux is constant. Results of dynamic simulation are presented in Fig. 11.23. Starting from

Solketal

428

Chapter 11 FUEL ADDITIVES

Table 11.11 Units Sizing and Economic Evaluation. Reactor Temperature/(
C) Pressure/(bar) Volume/(m3) Flow/(kg/h) Installed cost/(103 US$) Separation columns Number of trays Reflux ratio Distillate:Feed ratio Diameter/(m) Reboiler duty/(kW) Condenser duty/(kW) Installed cost/(103 US$) Utilities/(103 US$/year) Heat exchanger Temperature in/(
C) Temperature out/(
C) Pressure/(bar) Heat duty/(kW) Installed cost/(103 US$) Utilities/(103 US$/year) Mixers Temperature/(
C) Pressure/(bar) Volume flow/(m3/h) Installed cost/(103 US$)

Reactor 35 2 4 2049 81.04 COL-1 23 4 0.78 0.9 1227 1172 553.05 264.13

COL-2 11 2 0.434 0.3 137 110 116.27 28.23

COL-3 6 0.169 0.765 0.45 38 67 75.01 10.10 HEATER 54.1 35 2 26 14.98 15.67

MIX-A 53.5 1 2.197 23.46

MIX-G 68.9 0.064 0.313 6.97

steady state, the production rate is increased (top diagrams) or decreased (bottom diagrams) by changing the plant-inlet flow rate of glycerol (G-0), from the initial value of 3.2 kmol/h to 4 kmol/h and 2.4 kmol/h, respectively. The new production rate is reached in about 8 h. As more glycerol is fed to the plant, the flow rate of fresh acetone (A-0) is adjusted to the stoichiometric value, although the duration of the transient regime is quite long. The purity of solketal and water products exceeds 0.998. It should be remarked that other control strategies work equally well. For example, one can keep constant the reactor-inlet flow rate of acetone (instead of flow ratio). When the flow rate of fresh

429

Chapter 11 FUEL ADDITIVES

Table 11.12 Stream Table. Stream Name

G-0

A-0

A-REC

G-REC

RIN

Temperature/(o C) Pressure/(bar) Mole flow/(kmol/hr) Mass flow/(kg/hr) Mole fraction Glycerol Acetone Water Solketal Ethanol

25 1 3.22 296.2

25 1 3.22 186.8

57 1 26.87 1474.5

194 0.064 0.99 91.0

35 2 34.29 2048.7

1 0 0

0 1 0

0 0.81 0.03

1 0 0

012 0.73 0.03

0 ROUT 35 2 34.29 2048.7

0 WSG 128 1.24 7.42 574.2

0 0.16 SG 199.5 1.12 4.20 516.1

0 W 99 1 3.22 58.1

0.12 S 107 0.05 3.22 425.1

0.03 0.64 0.12 0.09 0.12

0.13 0 0.43 0.43 0

0.23 0 0 0.77 0

0 0 1 0 0

0 0 0 1 0

Temperature/(o C) Pressure/(bar) Mole flow/(kmol/hr) Mass flow/(kg/hr) Mole fraction Glycerol Acetone Water Solketal Ethanol

Figure 11.22 Plantwide control scheme for ketalization of glycerol with acetone.

430

Chapter 11 FUEL ADDITIVES

Figure 11.23 Solketal plantedynamic simulation results. Topdproduction rate increase by 25%; bottomdproduction rate decrease by 25%.

glycerol is available as manipulated variable, one can switch the flow and level control loops around the glycerol mixer.

11.5 Glycerol Acetalization Acetals are obtained from the condensation reaction of alcohols and aldehydes. Acetals with low molecular weight are used in cosmetics, food, or pharmaceutical industry. Acetals with high molecular weight are suitable diesel additives (Silva et al., 2010; Agirre et al., 2013). Conventional acetalization processes take place in homogeneous catalysis, using strong acids such as H2SO4, HF, HCl, or p-TSA (Agirre et al., 2013). This leads to corrosion problems and is not environmentally friendly, which has an important impact on the economic indicators. These problems could be avoided by performing the reaction in the presence of a solid catalyst, such as Amberlyst-15 (Delfort et al., 2005; Silva et al., 2010), Amberlyst 36 (Deutsch et al., 2007), and Amberlyst 47 (Agirre € emez et al., 2013). et al., 2011a; Gu

Chapter 11 FUEL ADDITIVES

(A) O OH +

HO

R

O

R

O

R'

OH

R'

+ 2 H2 O O

OH

R

HO

R'

O

(B)

O R' O

R' OH + R

HO

O

OH

O

OH

R

+ 2 ROH O

HO

R' O

Figure 11.24 Glycerol acetalization (A) and transacetalization (B).

Acetals of glycerol (Delfort et al., 2005) were proved to be valuable additives for diesel, biodiesel, or their mixtures (Rahmat et al., 2010). Thus, fuels containing oxygenated derivatives of glycerol have a positive impact on the economics of biodiesel production processes. Glycerol acetals can be obtained by reaction of glycerol with an aldehyde (direct acetalization) or by reaction of glycerol with another acetal (transacetalization) (Deutsch et al., 2007; Agirre et al., 2011a; Hong et al., 2013) (Fig. 11.24). In both cases, the reactions lead to two different isomers, with five or six heterocycles. The ratio between these isomers is about 1:1, but this depends on the temperature, the glycerol to aldehyde ratio, or the solvent (Behr et al., 2008).

11.5.1 Basis of Design In the following, the design of a process for manufacturing the glycerolebutanal acetals will be presented. The plant processes 452 kg/h aqueous solution of glycerol (mass fraction 0.695), the by-product of a biodiesel plant with the capacity of 3990 tonnes/year.

431

432

Chapter 11 FUEL ADDITIVES

11.5.2 Chemical Reaction € emez et al. (2013) performed batch experiments, in which the Gu reaction of glycerol (G) with butanal (B) took place in the presence of Amberlyst 47 as heterogeneous catalyst. The reaction leads to two isomers, namely dioxolane (5 atoms ring) and dioxane (6 atoms ring). During the reaction, the ratio of dioxane to dioxolane varies between 0.8 and 1.02. In the following, the isomers will be lumped in one component, generically denoted as “acetal” (A). The physical properties used during the simulation correspond to 5-hydroxy-2-alkyl-1,3-dioxane. This approximation is justified by the similar structure of the two isomers, which give similar properties estimated by group contribution methods. The reaction can be written in the following simplified form: k1

B þ G%A þ W k2

The reaction rate is given by r ¼ k1 $cB $cG  k2 $cA $cW

(11.16)

where the reaction rate constants follow an Arrhenius dependence on temperature Ea;j

kj ¼ kj0 e RT ; j ¼ 1; 2

(11.17)

€ emez et al., 2013) are The values of the kinetic parameters (Gu shown in Table 11.13. Equilibrium calculations show that, at temperatures between 50 and 100
C, high conversions are possible. Fig. 11.25 shows the equilibrium conversions of glycerol (XG) and butanal (XB) at different temperatures, calculated for several values of the initial glycerol/butanal ratio (MG,B). When butanal is in excess

Table 11.13 Glycerol Acetalization With Butanal. Kinetic Parameters. Parameter

Value

k01/(L2/mol$min$gcat) k02/(L2/mol$min$gcat) E1/(kJ/mol) E2/(kJ/mol)

5.33  105 267  1010 55.6 115

Chapter 11 FUEL ADDITIVES

1

1

0.2

0.8

3

0.8

1

1

0.6

XB

XG

0.6

433

MG,B = 3

0.4

0.4

0.2

MG,B = 0.2

0.2

0

0 50

60

70

80

90

100

50

60

T / [C]

70

80

90

100

T / [C]

Figure 11.25 Equilibrium conversion versus temperature, for different values of the glycerol/butanal ratio, MG,B. Continuous and dashed lines are values calculated for a water/glycerol ratio MW,G ¼ 0 and MW,G ¼ 1, respectively.

(MG,B ¼ 0.2), the glycerol conversion is almost complete. Similarly, an excess of glycerol (MG,B ¼ 3) leads to almost complete butanal conversion. The presence of water in the initial mixture (dashed line, water/glycerol molar ratio MW,G ¼ 1) does not have a significant effect on the chemical equilibrium. These results suggest that the reaction can be performed at an € emez excess of either reactant. However, the experiments (Gu et al., 2013) showed that, if glycerol is the limiting reactant, small amounts of 2,4,6-tripropyl-1,3,5-trioxane are formed by butanal cyclotrimerization. For this reason, the plant will be designed such that the reaction proceeds at an excess of glycerol. Thus, the reaction will take place at 70
C, with a molar ratio butanal/glycerol MB,G ¼ 1/3. The reactor will be sized such that the conversion of butanal (limiting reactant) is XB ¼ 0.75.

11.5.3 Conceptual Design The reactor-outlet stream contains unreacted glycerol and butanal, as well as the products acetal and water. In this mixture, the butanalewater azeotrope has the lowest boiling point (Table 11.14). Thus, a first distillation column (COL-1) will separate, as distillate, this azeotrope together with the rest of butanal, being recycled to the reactor. The bottoms stream contains water (W), glycerol (G), and acetal (A), from which the water can be easily separated in a second distillation column (COL-2). The last column (COL-3), operated under vacuum, performs the acetal/glycerol split. Because the difference in boiling points is quite large, obtaining high-purity acetal (above 99.5% molar) is easy. Moreover, getting glycerol with high purity is not necessary, this being recycled to the reaction section. Fig. 11.26 presents the plant flowsheet.

434

Chapter 11 FUEL ADDITIVES

Table 11.14 Boiling Points of Pure Components and Azeotropes (1 Atm). Component

Boiling Point

Binary azeotrope. (Butanal 0.7477 mol; water 0.2523 mol) Butanal Water Acetal Glycerol

69.5
C (unstable node) 74.8 (saddle) 100
C (saddle) 247.4
C (saddle) 287.7
C (stable node)

B-REC/W-REC B-0 W

LC PC

MIX-B

LC

FC

HEX

FC

TC

COL-1 REACTOR

LC

RIN

PC

TC

FC

LC

PC

FC

LC

A

ROUT TC

G-0

LC

COL-2

FC

TC

LC

MIX-G

COL-3

TC

LC

LC

G-REC

Figure 11.26 Glycerol acetalization plant. Flowsheet and control structure.

11.5.4 Reactor Design Because of the small amount of catalyst that is necessary (5% wt. catalyst/glycerol), the reaction will be performed in a stirred tank reactor. To account for external and internal resistance to mass transfer, we amended the kinetic constants by an effectiveness factor h ¼ 0.1. The catalyst can be retained inside the reactor by placing it in a basket and separation in a cyclone or an external decanting unit. A preliminary mass balance allows calculation of the reactor-inlet flow rates. Thus, for a conversion XB ¼ 0.75, the amount of catalyst is 56.7 kg. Considering a catalyst loading of 100 kg/m3, the reactor volume is about 0.6 m3.

Chapter 11 FUEL ADDITIVES

11.5.5 Separation Section The reactants/products separation can be achieved by distillation, as previously explained. The columns were designed using the RADFRAC model of Aspen Plus. The column diameter was found using the tray-sizing utility implemented in Aspen Plus. The simulation also gave the condenser and reboiler duties, which allowed sizing of the heat transfer equipment (a heat transfer coefficient of 500 W/m2/K was assumed). Table 11.15 Present the Main Equipment Sizes, Together With an Economic Evaluation. Detailed Stream Results are Given in Table 11.16.

Table 11.15 Equipment Sizing and Economic Evaluation. Reactor Temperature/(
C) Pressure/(bar) Volume/(m3) Mass flow rate/(kg/h) Catalyst mass/(kg) Installed cost/(103 US$) Distillation columns Number of trays Reflux ratio Distillate:feed ratio Diameter/(m) Reboiler duty/(kW) Condenser duty/(kW) Installed cost/(103 US$) Utility costs/(103 US$/an) Heat exchanger Inlet temperature/(
C) Outlet temperature/(
C) Pressure/(bar) Duty/(kW) Installed cost/(103 US$) Utility cost/(103 US$/an) Mixer Temperature/(
C) Pressure/(bar) Volumetric flow rate/(m3/h) Installed cost/(103 US$)

Reactor 72.8 1.2 0.5 1724.4 100 22.21 COL-1 12 5 0.059 0.21 159.3 89.6 109.83 30.52

COL-2 10 2 0.448 0.875 421.05 401.5 273.55 90.59

COL-3 19 5 0.252 0.758 236.7 224.5 247.09 50.86 HEX 105.7 72.8 1.2 66.9 27.69 4.03

MIX-1 36.8 1 0.42 8.37

MIX-2 145.1 1 1.2 16.1

435

436

Chapter 11 FUEL ADDITIVES

Table 11.16 Stream Table. Temperature/(o C) Pressure/(bar) Molar flow rate/(kmol/hr) Mass flow rate/(kg/hr) Molar fraction Glycerol Butanal Water Acetal Temperature/(o C) Pressure/(bar) Molar flow rate/(kmol/hr) Mass flow rate/(kg/hr) Molar fraction Glycerol Butanal Water Acetal

G-0

B-0

GW-1

BW-1

BW-REC

G-REC

25 1 11.06 452.1

25 1 3.41 245.71

145.1 1 21.22 1391.26

36.8 1 4.97 333.14

69.2 1 1.563 87.43

213.5 0.104 10.16 939.16

0.308 0 0.692 0 RIN 72.8 1.2 26.19 1724.4

0 1 0 0 ROUT 72.8 1.2 26.19 1724.4

0.636 0 0.361 0.003 WAG 134.6 1.11 24.63 1636.98

0 0.906 0.094 0 AG 181 0.077 13.58 1437.97

0 0.701 0.299 0 W 32.9 0.05 11.05 199.01

0.993 0 0 0.007 A 150.2 0.05 3.42 498.8

0.515 0.172 0.31 0.003

0.385 0.042 0.44 0.133

0.41 0 0.449 0.141

0.743 0 0 0.256

0 0 1 0

0.001 0 0.003 0.995

11.5.6 Process Control The process control structure is shown in Fig. 11.26. The glycerol flow rate is fixed by an upstream unit. At reactor inlet, the ratio between butanal (fresh and recycled, containing some water) and glycerol (aqueous solution) is fixed. The flow rate of fresh butanal is given by the level controller of the buffer vessel (MIX-B). The reaction temperature and reactor holdup are controlled by the duty and the outlet flow. Control of distillation columns is standard: pressure is controlled by the condenser duty; reflux drum and column sump levels are controlled by the distillate and the bottoms rates, respectively; one temperature in the stripping section is controlled by the reboiler duty; the reflux is constant. Results of dynamic simulation are shown in Fig. 11.27. Starting from steady state, the production rate is increased (A, B) or reduced (C, D), as a result of changing the fresh glycerol flow rate (G-0) by 25%, from the initial value 3.2 kmol/h to 4 kmol/h and 2.4 kmol/h, respectively. The new production rate is attained in about 10 h.

Chapter 11 FUEL ADDITIVES

437

Figure 11.27 Dynamic simulation results. The production rate is increased (A, B) or decreased (C, D) by 25%.

11.6 Concluding Remarks Glycerol, a versatile by-product of biodiesel production, can be converted into more valuable products by several reactions. This chapter uses published data concerning the reaction conditions (temperature, pressure, catalyst) and reaction rate to design processes in which glycerol reacts with isobutene, tert-butanol, butanal, and acetone to give ethers, acetals, or ketals which have various useful applications. The processes can be performed at industrial scale in simple RSR processes, using stirred tank reactors with the catalyst immobilized by different techniques, under relatively mild conditions and at reduced operating and investment costs. Thus, the total annual cost ranges from about 80 $/tonne (butanaleglycerol acetal) to 85 $/tonne (GTBE from isobutene) and 130 $/tonne (solketal). Etherification of glycerol with tert-butanol is more expensive (310 $/tonne), but this cost is greatly reduced to only 25 $/tonne by employing process intensification by RD technology. Thus, the main drive for the economics of these processes appears to be the cost of raw materials and the selling price of the products.

438

Chapter 11 FUEL ADDITIVES

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Chapter 11 FUEL ADDITIVES

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