Fuel 111 (2013) 449–460

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Fuel

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1-Butene oligomerization over ZSM-5 zeolite: Part 1 – Effect of reactionconditions

0016-2361/$ - see front matter � 2013 Elsevier Ltd. All rights reserved.http://dx.doi.org/10.1016/j.fuel.2013.03.066

⇑ Corresponding author. Tel.: +351 21 841 78 90; fax: +351 21 841 91 98.E-mail address: mandal@ist.utl.pt (M.A.N.D.A. Lemos).

A. Coelho a, G. Caeiro b, M.A.N.D.A. Lemos a,⇑, F. Lemos a, F. Ramôa Ribeiro a

a IBB – Institute for Biotechnology and Bioengineering, Centre for Biological and Chemical Engineering, Instituto Superior Técnico, UTL, Av. Rovisco Pais, 1049-001 Lisboa, Portugalb Galp Energia, SGPS, S.A., Rua Tomás da Fonseca, 1600-209 Lisboa, Portugal

h i g h l i g h t s

� Systematic study of the conversion of1-butene into fuels byoligomerization over HZSM-5catalysts.� Optimum conditions for selectivity

for Cþ8 hydrocarbons (�86 wt.%) wasobtained with H-ZSM-5, at 200 �C and0.5 bar.� HZSM-5 is a suitable catalyst and

allows almost steady-state operation.� Deactivation, with consequent

decrease in acidity, improves theperformance of the catalyst undersome conditions.� Low temperature performance of the

catalyst seems to be hindered by thevolatility of the products.

g r a p h i c a l a b s t r a c t

a r t i c l e i n f o

Article history:Received 7 June 2012Received in revised form 24 March 2013Accepted 25 March 2013Available online 9 April 2013

Keywords:1-Butene oligomerizationH-ZSM-5 zeoliteCatalyst deactivationHeterogeneous catalysisLight olefins

a b s t r a c t

The demand for middle distillates (kerosene and diesel) in comparison to gasoline fractions is increasingconstantly, particularly in the majority of European countries. Thus, maximizing the production of middledistillates in the refining process is of immediate interest to refiners. However, the production of liquidfuels, for instances using a catalytic cracker, always generates a significant amount of lighter fractions, inparticular of olefinic nature. Thus, the use of oligomerization reactions to convert the lighter olefin cutsinto middle distillates to incorporate in the diesel pool is a promising process for the production of cleandiesel fractions. In this work, 1-butene oligomerization over H-ZSM-5 zeolite has been investigated in adifferential reactor operating at ambient pressure. The effect of the reaction conditions, such as reactiontemperature, contact time and partial pressure was studied on the activity, selectivity and stability of thecatalyst. The results show that an increase in the reaction temperature and/or partial pressure and in thecontact time produces an improved catalyst activity. The data also show that high partial pressureimproves the selectivity to Cþ8 . Moreover, when increasing the temperature from 150 �C to 200 �C theCþ8 hydrocarbons selectivity increases, whereas above this temperature it decreases as expected, due tothe competition of cracking reactions. Furthermore, a decrease in contact time between the reaction mix-ture and the acid sites of the catalyst caused the Cþ8 hydrocarbons fraction in the product to increase. Thehighest selectivity towards Cþ8 hydrocarbons (�86 wt.%) was obtained with H-ZSM-5, at 200 �C and50 kPa of partial pressure and for the lowest value of contact time analyzed (12.5 � 10�3 h). In thisway, the operating conditions must be tuned in order to achieve a significant conversion and selectivityin desired fraction (Cþ8 ).

� 2013 Elsevier Ltd. All rights reserved.

450 A. Coelho et al. / Fuel 111 (2013) 449–460

1. Introduction

Olefin reactivity over acid catalysts (mainly zeolites) has beenextensively studied, in order to understand the mechanism of cat-alytic cracking which has been explained in terms of a carbeniumion mechanism [1–9]. However, the oligomerization reaction, ta-ken as the reverse of the cracking, is also a relevant process andit has also been receiving considerable attention by researchers[10–16] since it provides a way to convert surplus lighter olefins(such as propene and butene fractions which are generally ob-tained from the FCC-Fluid Catalytic Cracking effluent) into heavierhydrocarbons which are useful as sulfur-free fuels, mainly gasoline(C5AC10) [13,14].

At present, from an industrial viewpoint, the most widely usedprocess for producing gasoline from oligomerization of light olefinsis the Catpoly process developed by Universal Oil Products Com-pany (UOP) in the 1930s, which uses the phosphoric acid sup-ported over Kieselguhr silica clay (SPA) [17–19] Despite, theirextensive commercial application, this catalyst has several disad-vantages, namely the catalyst life time is relatively short, there isalso environmental concern associated with its disposal and corro-sion problems, and there is no possibility of reusing the spent cat-alyst. Furthermore, the products obtained are nearly exclusively inthe gasoline fractions [17,20].

In this way, the search for a possible replacement of phosphoricacid with other acid catalysts, easier to use and more environmen-tally friendly, led to the emergence of many different processes. Alot of development has been made in the field of homogeneous cat-alysts [21] and, in the past decades, a significant research effort hasbeen also directed towards the development of heterogeneous cat-alysts [19,22,23]. From a practical point of view, the latter optionshould always be the first choice both because of the ease of thecatalyst separation and due to their lower environmental impact[19,24].

Among the various processes developed are the Selectopol andPolynaphta processes (from IFP/Axens) which are based in silicaalumina amorphous [25,26], Shell’s SGPK process based in theNi-mordenite catalyst [27–29] and one of the most referenced pro-cesses, which dates from the 1980s, MOGD process-Mobil olefin togasoline and distillate developed by Mobil Research and Develop-ment Corporation [30–33]. This latter process uses ZSM-5 zeolite,to convert light olefins (C3AC4) from Fluid Catalytic Cracking(FCC) into higher molecular weight products such as gasoline ordiesel fuels. This process assumes special importance in the pres-ent since most of other processes were often oriented towardsthe production of gasoline [34] or to produce high-octane gasolineblending components [29]. However, as mentioned earlier, re-cently, the fuel market is shifting towards an increase in diesel con-sumption in detriment of gasoline, especially in the majority ofEuropean countries [35]. At the same time, environmental legisla-tion has led to more strict specifications on petroleum derivates[36]. In this way, the refining industry needs to adjust their pro-duction scheme in order to meet both the growth of the overallmarket and the shifting of the demand as efficiently as possible.ZSM-5 is a promising candidate for diesel production from olefinoligomerization and more studies using this zeolite seemed to beappropriate.

ZSM-5 zeolite (MFI topology) is a well-known crystalline micro-porous aluminosilicate which possess a tree-dimensional structurewith both straight and sinusoidal pores and it may have strongacidity with acid sites with different strengths [37]. The shapeselectivity properties of the ZSM-5 favor the formation of linearhydrocarbons [31,37,38]; this is particularly suitable for the pro-duction of good quality distillate fuels (diesel) [39]. Moreover, thiszeolite exhibits slower deactivation rates than other zeolites with

similar acidity [9,10,40]. For these reasons, ZSM-5 zeolite has beenwidely used in a variety of industrial processes. For example, ZSM-5 is used as an additive in the FCC catalyst to improve octane num-ber of FCC naphtha [1,27,41], it is also used in the disproportion-ation of toluene, xylenes isomerization, alkylation [42] andaromatization reactions [8,9]. It has also been found to be particu-larly suitable for converting methanol to gasoline (MTG) and ole-fins (MTO) [7–9].

It is widely accepted that the oligomerization reaction of olefinsover acid catalysts occurs via a carbocation mechanism[10,11,38,43] and its activity and selectivity strongly depend onthe operating conditions, through thermodynamics, kinetic, diffu-sional limitations and shape-selectivity effects [12,38]. In thisway, an understanding of the fundamental aspects of this reactionis still needed. With this in mind, in the present work, we investi-gated the effect of operating conditions, such as reaction tempera-ture, partial pressure and contact time between the reactionmixture and catalyst acid sites on the activity, selectivity and prod-uct distribution as well as on the stability of H-ZSM-5 zeolite cat-alyst. This work has been focused on a range of temperatureswhich favors oligomerization, which is lower than the one usedin other studies on the transformation of light olefins over ZSM-5, typically in a range of temperatures more related to catalyticcracking [44].

2. Experimental

2.1. Materials

Prior to use, the zeolite NH4-ZSM-5 (Zeolyst, Si/Al = 15, CBV3024E) was calcined in a tubular reactor at 500 �C for 8 h under aflow of dry air of 0.5 Ldry air g�1

zeo h�1, in order to remove any organiccompounds that could be adsorbed on the zeolite and to obtain itsprotonic form (H-form). After calcination, the zeolite was alwaysmaintained in a container with a constant and high humidity in or-der to ensure that the surface of the zeolite is fully saturated, pro-tecting it from adsorbing any contaminant, and at the same time,to keep the amount of water adsorbed on the catalyst constant.

1-Butene and the other gases used were high-purity products:1-butene (AirLiquide, 99%) nitrogen (AirLiquide, 99+%), air (Air-Liquide, 99+%), hydrogen (AirLiquide, 99+%) and ammonia (Linde,99.5%) and were dried by molecular sieves before use.

2.2. Catalytic test

Prior to each experiment, the catalyst was pre-treated under aflow of dry nitrogen of 60 mL min�1, to clean the surface of the cat-alyst, by heating it at 10 �C min�1 up to the temperature of 450 �C,where it was keep for 8 h. The reactor was then allowed to cooldown to the reaction temperature under the flow of N2.

The catalytic tests were performed using 1-butene as reactant,diluted with dry nitrogen in a fixed bed continuous flow reactor,in the gas phase, at atmospheric pressure and within a tempera-ture range from 150 �C to 250 �C. The partial pressure of the reac-tant was modified, from 12.5 kPa to 50 kPa, by changing thenitrogen flow for the same olefin flow (25 mL min�1), which resultsin the same contact time between the olefin and the catalyst. Theamount of H-ZSM-5 catalyst used in these tests was always50 mg. However, at 200 �C and with a partial pressure of 50 kPasome tests with 150 mg and 300 mg of the catalyst were also car-ried out, keeping the flow rates constant, in order to study the ef-fect of contact time. The reaction conditions used in this work aregiven in Table 1.

To avoid the deactivation of the catalyst that would occur dur-ing the chromatographic analysis, the reaction was interrupted by

Table 1Operating conditions for the catalytic transformation of 1-butene.

Reactiontemperature(�C)

Partialpressure(kPa)

Total flow (nitrogenplus olefin)(mL min�1)

Catalystmass(mg)

Contacttimea 1/WHSV (h)

150 12.5 200 50 12.5 � 10�3

200 25 100225 50 50250200 50 50 50 12.5 � 10�3

150 37.1 � 10�3

300 74.2 � 10�3

a Contact time values were determined with the mass of dry catalyst.(WHSV = weight hourly space velocity).

A. Coelho et al. / Fuel 111 (2013) 449–460 451

cutting off the olefin feed. During these periods, the sample waskept only under a flow of pure nitrogen at the working tempera-ture. Thus, the experiment consists of a sequence of feeding cyclesof the reactant, with 3 min each at the end of which the outletstream is sampled into the gas chromatograph.

2.3. Analysis of the products

All products were analyzed on-line with a Shimadzu GC-9A gaschromatograph equipped with a PLOT (KCl/Al2Cl3) capillary col-umn (50 m) and a Flame Ionization Detector (FID). The chromato-gram was integrated with a Shimadzu C-R3A integrator. Theproducts were lumped into several fractions based on their respec-tive GC retention times, identified by the injection of pure hydro-carbons calibration mixtures.

The activity, which measures the rate of consumption of the re-agent consumed per unit mass of dry catalyst, was defined accord-ingly to the following equation:

Activity ¼ F0xfinal

Wð1Þ

}where F0 is the molar flow rate of reagent (mol min�1), xfinal theconversion at the exit stream and W is the mass of dry catalyst (g).

Note that all the thermodynamic equilibrated butene isomerswere considered as reactants [45,46].

2.4. Catalyst physicochemical characterization

2.4.1. Elemental analysesThe concentration of silicon and aluminium in the zeolite was

obtained by elemental analyses. ICP (Inductively Coupled Plasma)was used to determine aluminium, while silicon was measuredby atomic emission spectrometry. The results of elemental analysisare presented in Table 2.

2.4.2. Scanning electron microscopyScanning electron microscopy (SEM) was performed to evaluate

crystallite size and morphology of the H-ZSM-5 used. This tech-nique was carried out in a JEOL JSM-7001F Field Emission Gun(FEG-SEM).

Table 2Physico-chemical properties of H-ZSM-5 zeolite.

Al (%w/w)a Si (%w/w)a (Si/Al)molara Total aci

2.2 (0.81) 39 17.0 1.02

a From elemental chemical analysis. Value in mmol Al g�1 given in parenthesis.b From ammonia TPD experiments.c Vmicro, micropore volume, determined by the t-plot method.d Sext, external surface area, determined applying the t-plot method.

2.4.3. Nitrogen adsorptionNitrogen adsorption–desorption isotherms at �196 �C of the

calcined catalysts were obtained using a Micromeritics ASAP2010 apparatus to determine the micropore volume and externalsurface area.

The sample was previously degassed under vacuum at 150 �Cfor 24 h. The micropore volume (Vmicro) and the external surfacearea (Sext) were calculated using the t-plot method, with the Har-kins–Jura equation for determining the thickness of the adsorbedlayer. The nitrogen adsorption isotherm at �196 �C on the freshH-ZSM-5 sample is presented in Fig. 1a.

2.4.4. Ammonia temperature programed desorption (NH3-TPD)The acid properties of the catalyst were obtained by ammonia

temperature desorption (TPD). Previously, the zeolite was placedin a quartz reactor and it was pre-treated, during 8 h at 450 �Cunder a flow 60 mL min�1 of dry nitrogen. After cooling to90 �C, ammonia adsorption was performed by injection of thesuccessive pulses of ammonia, until the catalyst surface was sat-urated with ammonia. The amount of ammonia in the effluentHe stream was measured by a thermal conductivity detector(TCD).

The desorption experiments of the ammonia saturated zeolites,were carried out in the TA Instruments SDT 2960 simultaneousDSC–TGA apparatus using an alumina pan. The sample was heatedup to 700 �C with a heating rate of 10 �C min�1, and about 23 mg ofeach saturated catalyst were used. Prior to this desorption step, thesamples were outgassed under N2 flow (80 mL min�1) by heatingat a rate of 10 �C min�1 up to 150 �C, and this temperature wasmaintained for 30 min to remove physisorbed water and NH3.Blank experiments, where no ammonia was adsorbed, were alsoperformed to account for baseline drifts.

The TG–DSC apparatus was calibrated, according to the manu-facturer’s specifications, in relation to weight, temperature andDSC calibrations.

The acid strength distribution was obtained from the NH3-TPDexperiments by numerical deconvolution, as described elsewhere[47–49].

2.5. Analyses of the catalysts after reaction

The amount of coke, formed after 21 min of reaction, was deter-mined by TG/DSC analyses in a TA Instruments SDT 2960 simulta-neous DSC–TGA apparatus, under an air flow of 75 mL min�1;about 20–22 mg of coked catalyst were used. The sample is heatedup to 210 �C at 10 �C min�1; this temperature is maintained for30 min to remove any water that might be adsorbed in the zeolite.After this pre-treatment the burning of the coke is performed usingthe same air flow, by increasing the temperature, at 10 �C min�1,up to 800 �C; this temperature was maintained for an additional25 min, before cooling down to the room temperature.

Therefore, the amount of coke was estimated according to thefollowing formula:

Coke amount ðwt:%Þ ¼ mi �mf

mi� 100 ð2Þ

dity (mmol NH3 g�1)b Vmicro (cm3 g�1)c Sext (m2 g�1)d

0.186 50

0

40

80

120

160

200

0 0.2 0.4 0.6 0.8 1Relative pressure (P/P0)

V ads

(cm

3 /g) (

STP)

(a)

0.0E+00

5.0E-06

1.0E-05

1.5E-05

2.0E-05

2.5E-05

3.0E-05

150 250 350 450 550 650 750

T(ºC)

dq/d

t(mm

ol.s

-1) (I)

(IV) (III)

(II)

(b)

Fig. 1. H-ZSM-5 zeolite nitrogen adsorption isotherm at �196 �C (a), NH3-TPD curves for the H-ZSM-5 zeolite and (b) (I – simulated curve; II – experimental curve;IV – decomposition curves; and III – blank curve).

452 A. Coelho et al. / Fuel 111 (2013) 449–460

where mi represents the weight of coked catalyst after desorption ofwater and mf represents the weight of coked catalyst after burning-off the coke.

The catalyst deactivation was computed as:

Catalyst deactivation ¼ ai � af

ai� 100 ð3Þ

where ai is the initial activity determined for 3 min and af is theactivity determined after 21 min of reaction.

2.6. Testing for mass transport limitations

The existence of external mass transfer limitations was investi-gated experimentally for the catalyst H-ZSM-5 at 200 �C by mea-suring the conversion of 1-butene as function of the reactantflow, while maintaining the contact time between reactant andcatalyst constant, by increasing the catalyst mass. Three flow rates(25, 50 and 100 mL min�1) were tested with different amount ofcatalysts (50, 100 and 200 mg), respectively. Since the conversionvalues for all experiments were rather similar (8.0%, 8.6% and7.5%), indicating that the mass transport did not influence the reac-tion rate, it was concluded that there are no external diffusionlimitations.

Internal limitations were checked, using the Weisz–Prater crite-rion [50] and, as it has been reported elsewhere [49], there are nosignificant diffusion limitations.

3. Results and discussion

3.1. Physico-chemical characteristics of H-ZSM-5 zeolite

The chemical composition of the calcined material presented aSi/Al ratio of 17. The small difference, observed between this valueand the value provided by the manufacturer (Si/Al = 15) can be as-cribed to the differences in the analytical methods. The values inTable 2 show that the H-ZSM-5 zeolite presents an external surfacearea of 50 m2 g�1, and a micropore volume of 0.186 cm3 g�1 (seealso Fig. 1a) both values are in agreement to those reported in lit-erature [51]. Scanning electron microscopy (SEM) showed that theH-ZSM-5 presented crystallites of ca. 2 lm.

The acidity of the H-ZSM-5 zeolite was determined by NH3-TPD,as shown in Fig. 1b) and as expected it has acid sites with differentstrengths. Two desorption peaks are present, the low-temperaturepeak at 150–300 �C corresponds to relatively weak acid sites, whilethe high-temperature peak at 350–480 �C corresponds to strongeracid sites. A total amount of acid sites of 1.02 mmol g�1

zeo was

obtained for this zeolite (Table 2). This value is in good accordancewith the expected value based on aluminium content using the Si/Al ratio from the supplier (1.04 mmol g�1

zeo) although it is a littlehigher than the value obtained from elemental analysis reportedin Table 2 (0.81 mmol g�1

zeo), in line with the difference already dis-cussed on the Si/Al ratio obtained by elemental analysis.

3.2. Catalytic transformation of 1-butene

We will now present and discuss the main results obtained inthe transformation of 1-butene over the H-ZSM-5 catalyst usedin this work. The oligomerization reaction is a relatively straight-forward one: the reaction is started by the dimerization of two ole-fin molecules, forming a dimer which can then react with furthermonomers, to produce higher oligomers, or crack, producing light-er products. The global activity is mainly related to the ability ofthe catalyst to produce the first dimers; this reaction is bimolecu-lar, so it is expected to increase with pressure, and exothermic, sothe equilibrium displacement towards the products is expected tobe favored by lower temperatures; in contrast the secondary reac-tion, cracking, is the reverse of the oligomerization reaction andwill be, in principle, favored by higher temperatures. The main dif-ficulty in this reaction is, thus, to be able to perform the reaction ata rate as high as possible, which would imply using higher temper-atures, while keeping the selectivity towards the oligomers high,which will be disfavored by increasing temperatures [16].

It is also expected that some deactivation occurs, in particularby deposition of what is known as coke, on the surface of the cat-alyst. This deactivation, however, is expected to be very limitedsince the catalyst has relatively narrow pores that prevent cokeformation.

3.2.1. Effect of time-on-stream and contact timeFig. 2a presents some examples of the time-course evolution of

the reaction.As expected the catalyst suffers some deactivation and the con-

version decreases rapidly in the beginning and tends to stabilize.Also, as it would be expected, the conversion increases (althoughthe activity decreases somewhat – see Fig. 3) with increasing con-tact time.

Fig. 2b shows the evolution of Cþ8 selectivity with time-on-stream (t-o-s) and contact time. Again, as expected, the selectivitytends to decrease with an increase in contact time due to theextension of secondary reactions. Selectivity, however, shows a rel-atively lower sensitivity towards time-on stream – it tends to berelatively constant except for the lowest contact time where it

0

2

4

6

8

10

12

14

16

18

Time-on-stream (min)

1-B

uten

e co

nver

sion

(%)

1/ WHSV=0.0742 h1/ WHSV=0.0370 h1/ WHSV=0.0125 h

(a) 1/WHSV=74.2 x 10-3 h1/WHSV=37.1 x 10-3 h1/WHSV=12.5 x 10-3 h

0

10

20

30

40

50

60

70

80

90

100

0 3 6 9 12 15 18 21 0 3 6 9 12 15 18 21Time-on-stream (min)

Sele

ctiv

ity C

8+ (wt.

%)

1/ WHSV=0.0742 h1/ WHSV=0.0370 h1/ WHSV=0.0125 h

(b)

1/WHSV=74.2 x 10-3 h1/WHSV=37.1 x 10-3 h1/WHSV=12.5 x 10-3 h

Fig. 2. Effect of time-on-stream on the conversion of 1-butene (a) and selectivity to Cþ8 products (b) for 1-butene transformation over H-ZSM-5 zeolite for different contacttimes at 200 �C and 50 kPa of partial pressure.

A. Coelho et al. / Fuel 111 (2013) 449–460 453

increases slightly with time-on stream. This increase is probablydue to the fact that deactivation will occur first on the more acidicsites which are the ones that most favor cracking reactions [44,52–55].

The changes in selectivity are also depicted in Fig. 3b were it canalso be seen that the extent of oligomerization, as measured by theaverage number of carbon atoms in the products, steadily de-creases with increasing contact time (from 8.24 to 7.34 and 6.98as contact time increases from 12.5 � 10�3 h to 37.1 � 10�3 h and74.2 � 10�3 h). In Fig. 4a, a further detail on product distributionshows that the increase in contact time also decreases sharplythe formation of trimers, clearly favoring cracking products, Scrack-

ing (C3AC7) increases from 12 wt.% to 32 wt.% when increasing thecontact time from 12.5 � 10�3 h to 74.2 � 10�3 h. The aromaticsselectivity (mainly benzene, xylenes, ethylbenzene, toluene andC9 aromatics) was always less than 10 wt.%.

We will now look at the effect of contact time on deactiva-tion; deactivation is mainly due to the formation of coke, a gen-eral denomination to the non-desorbed products that preventaccess to the acid sites decreasing the number and average acid-ity of the accessible sites [56,57]. In Table 3 we can see that,although coking is usually associated with secondary reactions,coke deposition decreases with increasing contact time. Thismay indicate that the deactivation is associated with the forma-tion of the primary products, which also decrease with increas-ing contact time

0

2

4

6

8

10

12

14

16

18

1/ WHSV (h)

Con

vers

ion

(%)

0.0000

0.0002

0.0004

0.0006

0.0008

0.0010

0.0012

0.0014

0.0016

0.0018

0.0020

Act

ivity

(mol

.min

-1.g

-1)

(a)

0.00 0.02 0.04 0.06 0.08

Fig. 3. Effect of contact time (reciprocal of WHSV) for the initial (t-o-s = 3 min) on (a) coselectivity to Cþ8 (open symbols), and average carbon number (closed symbols) at 200 �C

This interpretation is also consistent with the results obtainedfor the coke deposited on the catalysts, which are shown inFig. 4b. We can observe that, apart from a water desorption peak(not shown) that occurs around 100 �C, the TG analysis of thehydrocarbons deposited on the catalyst shows two main peaks,one around 300 �C and another one around 500 �C. The first oneis assigned mostly to the desorption of heavy products (reversiblecoke), probably in the range from C12 to C20 in view of the evapo-ration temperature, and is associated with an endothermic DSCsignal, while the second one, which can be attributed to the burn-ing of heavier molecules, is associated with an exothermic DSC sig-nal. From these results we can see that, although the amount ofheavier coke is mostly unaffected by the increase in contact time,there is a clear decrease in the first one. This was further verifiedby the observation that, when performing the reaction at lowertemperatures, if the used catalyst was washed with a light solvent(such as hexane) at the end of the reaction a significant amount ofoligomers that were retained inside the porous system wererecovered.

3.2.2. Effect of reaction temperature and partial pressure1-Butene conversions and initial activities obtained as a func-

tion of reaction temperatures and partial pressures are depictedin Fig. 5, accordingly to the operating conditions listed in Table 1.The initial activities were determined using the conversion valuesobtained for the t-o-s = 3 min.

50

55

60

65

70

75

80

85

90

95

100

0.00 0.02 0.04 0.06 0.08

1/ WHSV (h)

Sele

ctiv

ity C

8+ (wt %

)

6.8

7.0

7.2

7.4

7.6

7.8

8.0

8.2

8.4

Ave

rage

Car

bon

Num

ber

(b)

nversion of 1-butene (open symbols) and catalyst activity (closed symbols) (b) andand for 50 kPa of partial pressure.

0

10

20

30

40

5060

70

80

90

100

C8+

crack

ing

dimers

trimers

C9-C11

Aromati

cs

Prod

uct s

elec

tiviti

es (w

t %)

1/ WHSV=0.0125 h1/ WHSV=0.0370 h1/ WHSV=0.0741 h

1/WHSV=12.5 x 10 -3 h1/WHSV=37.1 x 10-3 h1/WHSV=74.2 x 10-3 h

(a)

0.00

0.02

0.04

0.06

0.08

0.10

0.12

0.14

0.16

0.18

200 400 600 800T (ºC)

DTG

(mg

min

-1)

200 ºC12.5 x 10-3 h

37.1 x10-3 h

74.2x10-3 h

(b)

Fig. 4. Effect of contact time on (a) the selectivities in cracking (C3AC7), trimers (C12), C9AC11 products and dimers (C8) and on (b) coke deposited on the catalyst for 1-butenetransformation at atmospheric pressure over H-ZSM-5 zeolite at 200 �C and 50 kPa.

Table 3Coke content obtained from TG/DSC experiments and deactivation percentage, for 1-butene conversion over H-ZSM-5 at 200 �C and 50 kPa for three different contacttimes.

Contact time (1/WHSV) (h) Deactivation (wt.%) Coke (wt.%)

12.5 � 10�3 52.48 7.8337.1 � 10�3 40.80 7.5874.2 � 10�3 38.81 6.70

454 A. Coelho et al. / Fuel 111 (2013) 449–460

The results clearly show that 1-butene conversion is dependenton the reaction temperature and partial pressure of the olefin. Asexpected, by thermodynamic and kinetic considerations, 1-buteneconversion increases when temperature and/or partial pressure in-creases. Note that the first step of this transformation is the oligo-merization reaction, a second order exothermic reaction whichconsequently is thermodynamically favored at low temperature.Nevertheless, as temperature increases the reaction rate also in-creases, although the cracking reaction rate has a larger increase[37].

The influence of the time-on-stream (t-o-s) can also be seen inFig. 5b and as seen above, the conversion decreases gradually withreaction time. The results of the coke content deposited on the cat-alysts and deactivation %, at the end of each 1-butene experiment,

Fig. 5. Effect of partial pressure or reaction temperature on the initial activity (after(N 250 �C, r 225 �C, s 200 �C and j 150 �C) (b) on the 1-butene conversion over H-ZSM

after 21 min of time-on-stream, are shown in Fig. 5. It is interestingto note that, in all cases, the deactivation occurs from early valuesof time-on-stream (t-o-s) and that the catalyst usually tends to aquasi-steady-state activity after a period of initial deactivation,which can be useful for a practical point of view.

It was also observed, that the amount of the coke deposited onthe zeolites has a general tendency to increase with partial pres-sure (see Fig. 6a), in line with the changes in activity. This is alsoprobably due to the production of higher molecular weight oligo-mers inside the zeolite pores being fully consistent with the resultsobtained for selectivity to Cþ8 products. The influence of tempera-ture is less clear: the amount of coke deposited increases up to225 �C and then decreases when the temperature increases up to250 �C. It is likely that at lower temperatures the catalysts deacti-vation is mainly due to the formation of heavy oligomers that donot evaporate (which is consistent with the results presentedabove – see Fig. 4b) and at sufficiently high temperature (above225 �C), these heavy products can evaporate or be cracked intosmaller molecules that can be diffuse more easily out of the cata-lyst pores, reducing coke deposition.

However, to explain the fact that, the deactivation of the cata-lyst decreased or remain practically constant when there is an in-crease in the coke percentage (see Fig. 6), it is necessary to considerthat, the deactivation of zeolite catalysts by coke is caused either

3 min of reaction) (a) and influence of time-on-stream, for various temperatures-5 zeolite.

0

1

2

3

4

5

6

7

8

9

10

Temperature (ºC)

Cok

e (w

t %)

(a)

0

10

20

30

40

50

60

70

80

90

100

150 175 200 225 250 150 175 200 225 250Temperature (ºC)

Dea

ctiv

atio

n (%

)

(b)

Fig. 6. Coke content obtained from TG/DSC experiments (a) and deactivation percentage (b), for 1-butene conversion over H-ZSM-5 zeolite at (N) 50 kPa, (j) 25 kPa and(r) 12.5 kPa.

A. Coelho et al. / Fuel 111 (2013) 449–460 455

by poisoning of acid sites or by pore blockage. In the first case, onecoke molecule blocks one active site, in the second case, one cokemolecule blocking the access of reactants to, on average, more thanone active site of the catalysts [56–58]. Therefore, the deactivatingeffect is much more pronounced in the case of pore blockage. Bothdeactivation modes can occur simultaneously although one of thetwo mechanisms is usually predominant.

The influence of temperature and partial pressure on selectivityis depicted in Fig. 7.

0

10

20

30

40

50

60

70

80

90

100

Temperature (ºC)

Sele

ctiv

ity C

8+ (wt %

)

(a)

0

2

4

6

8

10

12

14

150 175 200 225 250

0 5 101-Butene co

Yiel

d C

8+ (w

t %)

(c)

Fig. 7. Effect of temperature or partial pressure ((N) 50 kPa, (j) 25 kPa and (r) 12.5 kPa(t-o-s = 3 min) (c), for the 1-butene transformation over H-ZSM-5 zeolite.

The results indicate that the selectivity towards Cþ8 increaseswhen the partial pressure of the reagent in the feed was increasedfrom 12.5 kPa to 50 kPa while keeping the temperature constant.As discussed above this result can be attributed to the fact thatoligomerization reaction is a second order reaction. In this way,although the oligomerization reaction can be conducted at a widerange of olefin partial pressures, higher olefin partial pressures arepreferred since they promote the bi-molecular reactions leading tooligomerization [59,60]. This effect is more noticeable at lowertemperatures.

0

2

4

6

8

10

12

14

Temperature (ºC)

Yiel

d C

8+ (w

t %)

(b)

125 150 175 200 225 250

15 20 25 30nversion (%)

) on the selectivity in Cþ8 (a) or yield in Cþ8 (b) and yield in Cþ8 vs. initial conversion

456 A. Coelho et al. / Fuel 111 (2013) 449–460

Changes in the reaction temperature have a more complex ef-fect in the selectivity in Cþ8 . Fig. 7a shows that the temperature risehas two distinct effects; it was found that by increasing the tem-perature from 150 �C to 200 �C the Cþ8 hydrocarbons in the productincreased, whereas above this latter temperature they decreased.The selectivity to the oligomerized products (Cþ8 ) goes through amaximum of approximately 86 wt.% at 200 �C and 50 kPa of partialpressure. The decrease of selectivity for higher temperatures wasexpected and can be explained by the fact that cracking reactions,having higher activation energy, are favored by the increasing intemperature.

However, the appearance of a maximum is not an expectedbehavior if one takes into account the kinetic and thermody-namic aspects of the reaction. Since the Cþ8 products result fromthe initial oligomerization reaction while products with a num-ber of carbon atoms lower than eight carbon atoms result fromcracking reactions that occur on these initial oligomers, it wouldbe expected that at low temperatures the Cþ8 selectivity shouldbe higher and decreased steadily with temperature. However,the experimental observation of a low selectivity for oligomeri-zation products at the lowest temperature studied (at 150 �C),can be explained by the fact that some of the oligomerizationproducts formed have boiling points above the operating tem-perature. Thus, they may be retained (through condensation/trapping) inside the pores of the zeolites, staying longer in con-tact with acid sites of the catalyst given the greater difficulty inleaving the pores owing to the strong adsorption [8] of theseproducts on the acid sites. For example, Cþ9 hydrocarbons have

0

10

20

30

40

50

60

70

cracking dimers trimers C9-C11

Prod

uct S

elec

tiviti

es (

wt %

)

150 ºC

200 ºC

225 ºC

250 ºC

(a)50 kPa

(

0

10

20

30

40

50

60

70

cracking dimers trimers C9-C11

Prod

uct S

elec

tiviti

es (

wt %

)

(c)200 ºC ■ 50 kPa

■ 25 kPa

■ 12.5 kPa

(

Fig. 8. Product distribution (wt.%) for 1-butene oligomerization at atmospheric pressure:on-stream at 200 �C and partial pressure 50 kPa (b) and effect of partial pressure at 200

boiling point temperatures above 150 �C [61]. This leads to anincrease of the probability of successive cracking reactions andconsequently to an increase in the cracking products (see alsoFig. 8a), consistent with the observed selectivity towards crack-ing products when the contact time between catalyst and prod-uct mixture was increased (see discussion above, in particular inrelation to Fig. 3).

The results also show that the Cþ8 yield increases both with reac-tion temperature and partial pressure although when workingwith lower temperatures or lower butene conversions the Cþ8 yieldis unaffected by partial pressure (see Fig. 7b and c).

These results indicate that when in practical operation of a reac-tor a compromise between selectivity and activity has to be con-sidered, and if the goal is to achieve a good selectivity on theoligomers production, lower activity values are certainly to beexpected.

A more detailed view of the influence of temperature on theproduct selectivities obtained for 1-butene transformation at50 kPa of partial pressure is shown in Figs. 8a and 9.

It can be seen that, with exception of 150 �C, the selectivity to-ward light hydrocarbons, C3AC7, (Scracking), increases with reactiontemperature while, the dimers product-C8 hydrocarbons (Sdimers),trimers product-C12 hydrocarbons (Strimers) and C9AC11 products(SC9AC11) selectivities show a general trend to decrease with tem-perature, in accordance with the fact that the lighter hydrocarbons(<C8) are the result of the cracking of the heavier products formedby the oligomerization reaction. These results indicate that at200 �C, oligomerization of olefins is the dominant reaction while

0

10

20

30

40

50

60

70

cracking dimers trimers C9-C11

Prod

uct S

elec

tiviti

es (w

t %)

t-o-s=3 min

t-o-s=9 min

t-o-s=21 min

b) 200 ºC; 50 kPa

0

10

20

30

40

50

60

70

cracking dimers trimers C9-C11

Prod

uct S

elec

tiviti

es (

wt %

)

d)250 ºC

■ 50 kPa

■ 25 kPa

■ 12.5 kPa

effect of temperature for t-o-s = 3 min and partial pressure 50 kPa (a), effect of time-�C and 250 �C (t-o-s = 3 min).

A. Coelho et al. / Fuel 111 (2013) 449–460 457

at higher temperature cracking is predominant leading to a highercracking selectivity (C3AC7 hydrocarbons). As discussed above, itwould be expected that at 150 �C the selectivity towards crackingproducts would be lower than those obtained at other tempera-tures studied, but this is not the case.

Partial pressure, on the other hand, has a much smaller influ-ence on the selectivities (see Fig. 8c and d), although an increasein dimers selectivity can be observed upon increase of the partialpressure, as it would be expected. In general we can say that, asthe partial pressure is increased, the selectivity towards crackingproducts decreases while the selectivity towards hydrocarbonswith eight or more carbon atoms increases.

Another interesting aspect is the change of the selectivity withtime-on-stream, depicted in Fig. 8b) for 200 �C and 50 kPa. Whilethe selectivity to cracking products does not change significantlythere is an enhancement of selectivity for the dimers, accompaniedby a decrease in the selectivity to trimers and C9AC11 products. Infact these observations are to be expected since secondary prod-ucts are the most affected with the decrease of conversion due tocatalyst deactivation.

Additional information on product distribution, includingparameters like olefins/paraffins, paraffins/aromatics, oligomeriza-tion/cracking, aromatization/cracking and aromatization/oligo-merization ratios, at different temperatures and partial pressuresare depicted in Figs. 9 and 10.

It was observed that substantial amounts of paraffins areformed, resulting from hydrogen transfer reactions involving reac-

0

5

10

15

20

25

30

35

Number of Carbon Atoms

Prod

uct D

istr

ibut

ion

(wt %

)

Paraffins

Olefins

Aromatics

150 ºC(a)

0

5

10

15

20

25

30

35

Number of Carbon Atoms

Prod

uct D

istr

ibut

ion

( wt %

)

Paraffins

Olefins

Aromatics

225 ºC(c)

1 2 3 4 5 6 7 8 9 10 11 12

1 2 3 4 5 6 7 8 9 10 11 12

Fig. 9. Product distribution for the transformation of 1-butene at differe

tant or product olefins. Under all the reaction conditions investi-gated, 80 wt.% or more of the products observed in the gas phasewere paraffins and olefins, in the range from C3 to C12. The remain-ing products (20 wt.% or less) were aromatics (mainly benzene, xy-lenes, ethylbenzene, toluene and C9 aromatics – see Table 3),probably formed by dehydrocyclization of the olefins.

The aromatics selectivity increases with increasing temperatureand partial pressure (except at 150 �C), as it can be seen in Table 4.This is in-line with what is thermodynamically expected [62] andin view of the fact that hydrogen transfer involves high activationenergy bimolecular reactions.

The aromatization/cracking ratio, however, was found to de-crease with increasing reaction temperature for the three partialpressures studied (Fig. 10a) whereas aromatization/oligomeriza-tion tends to increase with reaction temperature, except at150 �C. For a given temperature both ratios are higher at higherpartial pressures.

Higher aromatic selectivity observed at 150 �C (see Table 4)may be attributed to the same effect that was discussed above,where the low volatility of the oligomerization products inducesa larger contact time between the reaction mixture and the activesites of the catalyst.

In fact the thermodynamics for the secondary reactions (hydro-gen transfer, cracking and aromatization) indicates that all are fa-vored with increasing temperature [12].

The experimental olefin to paraffin ratio obtained as a functionof reactor temperature and olefin pressure is illustrated in

0

5

10

15

20

25

30

35

Number of Carbon Atoms

Prod

uct D

istr

ibut

ion

(wt %

)

Paraffins

OlefinsAromatics

(b) 200 ºC

0

5

10

15

20

25

30

35

431 2 5 6 7 8 9 10 11 12

1 2 53 4 6 7 8 9 10 11 12Number of Carbon Atoms

Prod

uct D

istr

ibut

ion

(wt %

)

Paraffins

Olefins

Aromatics

250 ºC(d)

nt reaction temperature at 50 kPa with the fresh H-ZSM-5 catalyst.

Fig. 10. Effect of temperature or partial pressure ((N) 50 kPa, (j) 25 kPa and (r) 12.5 kPa) on oligomerization/cracking (open symbols), aromatization/cracking (filledsymbols) and aromatization/oligomerization (grayed symbols) mass ratios (a) and for olefins/paraffins (b) and paraffins/aromatics (c) molar ratios for the 1-butenetransformation over fresh H-ZSM-5 zeolite.

Table 4Olefins/paraffins (O/P), paraffins/aromatics (P/A) C@

5 and C@3 olefins and hydrogen/

carbon (H/C) molar ratios obtained for the 1-butene transformation over H-ZSM-5zeolite at different temperatures and partial pressures of reactant.

Partialpressure (kPa)

Reactiontemperature(�C)

Aromatics(wt.%)

Molar ratio

O/P P/A

C@5 =C@

3 H/C

12.5 150 19.3 5.00 0.6 5.03 1.87200 4.9 1.55 8.8 3.39 2.06225 5.5 2.08 6.9 2.77 2.06250 5.9 3.69 5.1 2.05 2.04

25 150 15.4 2.88 1.5 4.87 1.93200 6.8 1.32 6.7 3.63 2.06225 7.3 1.90 5.4 2.78 2.05250 6.4 2.33 5.6 2.28 2.05

50 150 11.1 1.02 4.4 3.63 2.04200 9.1 0.73 6.2 3.93 2.07225 9.4 1.79 4.2 2.80 2.03250 10.8 2.25 3.3 1.80 2.01

458 A. Coelho et al. / Fuel 111 (2013) 449–460

Fig. 10b). It is possible to see that, in most cases and as expected,this ratio presents values greater than one [63,64]. This ratio isan indication of the extent of secondary reactions; in fact, shouldnone of these secondary hydrogen transfer reactions occur wewould expect that the olefin/paraffin ratio to be infinite.

It is possible to see, in Fig. 9, that increasing the temperaturefrom 200 �C to 250 �C clearly increases the production of C3 andC5 olefins, while C8 paraffins production decreases due to thecracking secondary reactions. However it can also be seen that C1

and C2 products (methane and ethane/ethylene, respectively) are

not observed, which is an indication that, as expected, the protolyt-ic cracking mechanism is negligible for these operating conditions.The same is true for the two other partial pressures studied (resultsnot shown).

A paraffin/aromatic ratio of three would be expected since forthe formation of an aromatic hydrocarbon from an olefin, threeother olefin molecules would have to be hydrogenated, producingthree paraffin molecules. However, in most cases, the molar paraf-fins/aromatics ratios have values greater than 3 and in some caseswith relatively high values (see Table 4 and also Fig. 10b). This factindicates that a significant part of the hydrogen responsible forparaffin production is associated with coke formation.

Also, it can be observed from Table 4 that, with few exceptions,the hydrogen/carbon ratio obtained in the products has a similarH/C ratio of the reactant (1-butene), presenting a value close to2. The values observed for H/C ratio are consistent with the paraffinto aromatic ratio discussed above and the fact that little coke is, infact, deposited on the catalyst.

The ratio between C@5 =C@

3 products was also calculated for allruns (see Table 4) since, apart from butane itself, these would bethe most probable products of dimer cracking. The results showedthat, for the same partial pressure, there is a general trend for theC@

5 =C@3 molar ratio to decrease with increasing temperature and al-

ways present values higher than 1, probably due to further trans-formation of the cracking products, which may be involveoligomerization reactions.

4. Conclusions

The results obtained in this systematic study showed that olefinoligomerization is a reaction for which the activity and selectivity

A. Coelho et al. / Fuel 111 (2013) 449–460 459

strongly depend on the operating conditions and where secondaryreactions, such as cracking, constitute a serious limitation. Temper-ature, reactant partial pressure and contact time between the reac-tion mixture and the acid sites of the catalyst have great effect incontrolling activity and selectivity of the catalysts. The resultsshowed an increase in the catalytic activity with increasing tem-perature, partial pressure and or contact time. However it wasfound that oligomerization selectivity is maximal at 200 �C. Athigher temperatures cracking increases rapidly, resulting in aproduct distribution that shifts towards lighter hydrocarbons.Although it would be reasonable to assume that at 150 �C theselectivity towards oligomerization should be still higher than at200 �C, in view of thermodynamic considerations, an unexpectedhigher cracking selectivity was obtained which was attributed tothe fact that some of the oligomerization products formed will re-main condensed/trapped within catalyst due to their low volatility.

It was also observed that the higher partial pressures favoredthe formation of heavier products while long contact time favorsthe cracking products. The results also show that the deactivation,which is accompanied by a lowering of the acidity of the catalyst,improved selectivity for dimer (C8) products.

According to the results previously described, the best condi-tions, in terms of the selectivity toward oligomerization products(Cþ8 ), when using this HZSM-5 (Si/Al = 15) zeolite, seem to be thecombination of higher partial pressures (50 kPa), lower value ofcontact time (12.5 � 10�3 h) and a temperature of 200 �C. Further-more, although the catalyst presents some initial deactivation, itsactivity stabilizes after a while and, under the right conditions, thislimited deactivation is accompanied by an increase in oligomeriza-tion selectivity. Under these conditions it was possible to obtainthe highest value of selectivity for Cþ8 (ca. 86 wt.%).

As a sequence for this work the influence of the acidity of thecatalyst should be investigated since it seems clear that very acidiccatalysts tend to favor cracking reactions in relation to oligomeri-zation. It can also be suggested that the oligomerization of olefinswould benefit from a process design that allowed the continuousremoval of the oligomerization products.

Acknowledgments

A. Coelho wishes to thank to Fundação para a Ciência e Tecno-logia for financial support for the PhD Grant (Ref. SFRH/BD/66744/2009). The authors also would like to thank José Roque (from GALPEnergia, Portugal) for fruitful discussions.

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