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Process simulation and catalyst development for biodiesel production 2006

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PROCESS SIMULATION AND CATALYST DEVELOPMENT FOR BIODIESEL PRODUCTION By Alex Harris West B.A.Sc, University of British Columbia, 2003 A THESIS SUBMITTED IN PARTIAL FULFILLMENT OF THE REQUIREMENTS FOR THE DEGREE OF MASTER OF APPLIED SCIENCE in The Faculty of Graduate Studies (Chemical and Biological Engineering) UNIVERSITY OF BRITISH COLUMBIA August 2006 © Alex Harris West, 2006 Abstract F o u r cont inuous b iod iese l processes were des igned and s imu la ted i n H Y S Y S . T h e f i rst t w o e m p l o y e d t rad i t ional homogeneous a l ka l i and ac id-cata lysts . T h e th i rd and four th processes used a heterogeneous ac id catalyst and a supercr i t ica l me thod , respect ive ly , to conver t a waste vegetable o i l feedstock into b iod iese l . W h i l e a l l processes were capab le o f p r o d u c i n g b iod iese l at h igh pur i ty , the heterogeneous and supercr i t ica l processes were the least c o m p l e x and had the smal lest number o f unit operat ions. M a t e r i a l and energy f l o w s , as w e l l as s i zed uni t operat ion b l o c k s , were used to conduc t an e c o n o m i c assessment o f each process. To ta l cap i ta l investment , total manufac tu r ing cost and after tax rate-of-return ( A T R O R ) were ca lcu la ted for each process. T h e heterogeneous ac id catalyst process had the lowest total cap i ta l investment and manufac tu r ing costs, and had the on l y pos i t i ve A T R O R . F o l l o w i n g the results o f the process s imu la t ions , tin(II) o x i d e was invest igated fo r use as a heterogeneous catalyst. Un fo r tuna te ly , cata ly t ic exper iments showed no act iv i ty . Subsequent ly , a carbon-based ac id catalyst was prepared by su l fonat ing py ro l ys i s char , and was studied fo r its ab i l i ty to ca ta lyze t ransester i f icat ion o f vegetable o i l . T h e catalyst showed on l y qual i ta t ive t ransester i f icat ion, but demonstrated good conve rs ion i n free fatty ac id ester i f ica t ion. Expe r imen ts were des igned to measure the effect o f a l coho l to o i l ( A : 0 ) mo la r rat io, react ion t ime and catalyst l oad ing on the sample . It was observed that f ree fatty ac id ( F F A ) conve rs ion increased w i t h inc reas ing A : 0 m o l a r rat io, react ion t ime and catalyst l oad ing . C o n d i t i o n s that y ie lded the greatest conve rs ion were 18:1 A : 0 mo la r rat io, 3 hour react ion t ime, 5 wt .% catalyst , 7 6 ° C under re f lux . T h e above cond i t ions reduced the F F A content i n a waste vegetable o i l ( W V O ) - e t h a n o l m ix tu re f r o m 4.25 wt .% to <0.5 wt.%. U n d e r an 78:1 A : 0 mo la r rat io and ident ica l cond i t ions , the catalyst was able to reduce the F F A content o f a W V O feedstock f r o m 12.25 wt .% to 1 wt.%. T h e catalyst has potent ia l to be used i n a process conver t ing a h igh F F A feedstock to b iod iese l i f the l im i ta t ions to t ransester i f icat ion can be ove rcome. O therw ise , it w i l l serve as an exce l len t catalyst for reduc ing the F F A content o f feedstocks in a two-step ac id and base conve rs ion process. i i Table of Contents Abs t rac t i i T a b l e o f Contents ; i i i L i s t o f Tab les v L i s t o f F igu res v i Nomenc la tu re v i i i A c k n o w l e d g e m e n t s i x 1 In t roduct ion 1 1.1 Transester i f i ca t ion research 2 1.1.1 H o m o g e n e o u s a l ka l i - ca ta l yzed t ransester i f icat ion 3 1.1.2 H o m o g e n e o u s ac id -ca ta lyzed t ransester i f icat ion 3 1.1.3 Hete rogeneous ly ca ta lyzed t ransester i f icat ion ; 4 1.1.4 Superc r i t i ca l t ransester i f icat ion , 5 1.2 Process m o d e l l i n g and e c o n o m i c assessment 6 1.3 Thes i s ob ject ives 6 1.4 T h e s i s format 7 1.5 References 8 2 Assessmen t o f F o u r C o n t i n u o u s B i o d i e s e l P r o d u c t i o n Processes us ing H Y S Y S . P l a n t 10 2.1 In t roduct ion and b a c k g r o u n d 10 2.2 Process s imu la t i on 13 2.3 Process des ign 15 2.4 E q u i p m e n t s i z i n g 16 2.4.1 Reac to r vessels 17 2.4.2 C o l u m n s 17 2.4.3 G r a v i t y separators 17 2.4.4 H y d r o c y c l o n e 18 2.5 E c o n o m i c assessment 18 2.5.1 B a s i s o f ca lcu la t ions , 18 2.5.2 T o t a l cap i ta l investment 19 2.5.3 To ta l manufac tu r ing cost 19 2.6 Sens i t i v i t y analyses and op t im iza t i on 21 2.7 C o n c l u s i o n 22 i i i References • 39 3 Charac te r i za t i on and Tes t i ng o f He te rogeneous Cata lys ts f o r B i o d i e s e l P r o d u c t i o n 41 3.1 In t roduct ion and backg round 41 3.2 T in( I I ) o x i d e synthesis and test ing methods 4 4 3.2.1 S n O synthesis procedure 4 4 3.2.2 Ca ta lys t test ing. •• 4 4 3.3 T in( I I ) o x i d e results and d i scuss ion 4 5 3.3.1 Synthes is and character iza t ion 45 3.3.2 Ca ta l y t i c ac t iv i ty ; •- 4 5 3.4 Su l fona ted char synthesis and test ing mthods 45 3.4.1 Su l fona ted char synthesis procedure 45 3.4.2 Su l fona ted char test ing procedure 46 3.5 Su l fona ted char results and d i scuss ion 47 3.5.1 Ca ta l ys t charac ter iza t ion 4 7 3.5.2 Su l fona ted char cata ly t ic ac t iv i ty 51 3.6 C o n c l u s i o n 5 4 3.7 References 66 4 C o n c l u s i o n , Gene ra l D i s c u s s i o n and R e c o m m e n d a t i o n s 68 4.1 G e n e r a l d i scuss ion '. 68 4.2 C o n c l u s i o n s 7 0 4 .3 R e c o m m e n d a t i o n s 7 2 4.4 References 74 List of Tables Table 1.1. Selected heterogeneous acid catalysts used for transesterification of triglycerides and their results , 5 Table 2.1. Catalysts and reaction parameters for heterogeneously catalyzed reactions of soybean oil at 1 atm 25 Table 2.2. Summary of unit operating conditions for each process 26 Table 2.3. Feed and product stream information for the alkali-catalyzed process 27 Table 2.4. Feed and product stream information for the homogeneous acid-catalyzed process. .27 Table 2.5. Feed and product stream information for the heterogeneous acid-catalyzed process. 28 Table 2.6. Feed and product stream information for the supercritical methanol process 28 Table 2.7. Equipment sizes for various process units in all processes. (Dimensions are diameter x height, m) 29 Table 2.8. Equipment costs, fixed capital costs and total capital investments for each process. (Units: $millions) 30 Table 2.9. Conditions for the economic assessment of each process. (Zhang et al. 2003b) 31 Table 2.10. Total manufacturing cost and after tax rate-of-return for each process. (Units: $millions) 32 Table 3.1. BET surface areas for each catalyst sample 47 Table 3.2. Mass per cent composition by element and molecular formula of each catalyst sample 48 Table 3.3. Total acidity for each catalyst sample 49 v List of Figures F igu re 2 .1 . H o m o g e n e o u s b a s e - c a t a l y z e d process f lowsheet (Process I) 33 F i g u r e 2.2. H o m o g e n e o u s a c i d - c a t a l y z e d process f lowsheet (Process II) 3 4 F i g u r e 2.3. Heterogeneous a c i d - c a t a l y z e d process f lowsheet (Process III) . . . .35 F i g u r e 2.4 Superc r i t i ca l a l coho l process f lowsheet (Process I V ) 36 F i g u r e 2.5. A f te r - tax rate o f return vs. react ion conve rs ion fo r a l l processes 37 F i g u r e 2.6. A T R O R vs . methanol recovery i n the methano l recovery c o l u m n , H A C process. . . 37 F i g u r e 2.7. A T R O R vs. operat ing pressure i n the methano l recovery c o l u m n , H A C process . . . . 38 F igu re 3 .1 . S a m p l e o f u n k n o w n substance obta ined dur ing S n O preparat ion v i a method o f A b r e u e t a l . (2005). . . . . 56 F i g u r e 3.2. C o m m e r c i a l sample o f S n O 56 F i g u r e 3.3. X R D pattern o f S n O sample prepared by method o f Fu j i t a et a l . (1990) 57 F i g u r e 3.4. X R D pattern o f c o m m e r c i a l S n O sample 57 F i g u r e 3.5. Ca ta lys t 1 X R D pattern 58 F i g u r e 3.6. X P S survey scan for Ca ta lys t 1 58 F i g u r e 3.7. N a r r o w scan in S 2p reg ion fo r Ca ta lys t 1 59 F i g u r e 3.8. N a r r o w scan i n C Is reg ion fo r Ca ta lys t 1 59 F i g u r e 3.9. n - P r o p y l a m i n e pu lse adsorp t ion peaks fo r Ca ta lys t 1 60 F igu re 3.10. T P D curve fo r Ca ta lys t 1. R a t i o o f weak ac id sites to strong ac id sites is 0 .85 :1 . . . 60 F i g u r e 3 .11 . T P D cu rve fo r Ca ta l ys t 2 . R a t i o o f weak a c i d sites to s t rong a c i d sites is 1.21:1.. . 61 F i g u r e 3.12 S E M image o f Cata lys t 1 ind ica t ing pore s izes 61 F i g u r e 3.13. S E M image o f Cata lys t 2 e m p h a s i z i n g f ib rous channels and pore ne twork 62 F i g u r e 3.14. S E M image o f Cata lys t 3, h igh l i gh t i ng var iab le s ize o f catalyst par t ic les 62 F igu re 3.15. E f fec t o f react ion t ime on f i na l ac id number . Reac t ions were run at 5 wt .% catalyst 1 w i th e thanol at A : 0 mo la r rat ios o f 6 : 1 , 9 .5 :1 , 18 :1 , 2 8 : 1 , 3 8 : 1 , 48:1 63 F igu re 3.16. E f fec t o f A : 0 mo la r rat io at f i x e d react ion t ime on final ac id number . 5 wt .% catalyst 1 63 F i g u r e 3.17. E f fec t o f A : 0 mo la r rat io on final ac id number fo r the 15 hour set o f react ions. 5 wt .% Cata lys t 1 64 F i g u r e 3.18. E f fec t o f catalyst amount o n final ac id number . 28:1 A : 0 mo la r rat io, e thano l , 5 wt .% catalyst 1 6 4 v i F i g u r e 3.19. F i n a l ac id number o f react ion m ix tu re after reac t ion w i t h each catalyst sample . 3 hour react ion , 28:1 A : 0 mo la r rat io, 5 wt .% cata lyst l oad ing 65 Nomenclature Definition Symbol Units After-tax rate of return ATROR - Alcohol to oil A:0 - Ammonia NH3 - Auxiliary facility cost CAC - Bare module capital costs CBM $ Bare module factor FBM - Bare module factor parameter B, - Bare module factor parameter B2 - Brunauer, Emmett and Teller BET - Capacity parameter A - Carbon dioxide C02 - Carboxylic acid group COOH7 - Contingency fee CcF $ Energy dispersive X-ray EDX - Fatty acid methyl-ester FAME - Fixed capital cost CFC $ Free fatty acid FFA - Gas chromatograph GC - Green house gas GHG - Heterogeneous Acid Catalyzed HAC - Materials factor FM - Non-random two liquid NRTL - Pressure factor FP - Purchase cost cP $ Purchase cost parameter K, - Purchase cost parameter K2 - Scanning electron microscopy SEM - Sulfate group S042' - Temperature programmed desorption TPD - Thermal conductivity detector TCD - Tin(II) chloride SnCh - Tin(U) oxide SnO - Total capital investment CTCI $ Total manufacturing cost TMC $ Total module cost CTM $ Waste vegetable oil wvo - Weight per cent wt.% - Working capital cost Cwc $ X-ray diffraction XRD - X-ray photospectroscopy XPS - viii Acknowledgements I a m ex t remely grateful to m y superv isor D r . N a o k o E l l i s , and commi t tee members D r . D u s k o Posarac and D r . J o h n R. G r a c e , for their support , gu idance and pat ience throughout the course o f th is degree. T h a n k you to D r . K e v i n J . S m i t h for the use o f h is laboratory fac i l i t ies , M r . Ib rah im A b u for h is assistance w i th the B E T measurements and D r . X u e b i n L i u fo r h is t remendous assistance w i t h the n -p ropy lamine exper iments and interpret ing the resul ts. T h e f i nanc ia l support o f the Natu ra l Sc iences and E n g i n e e r i n g Resea rch C o u n c i l is gratefu l ly acknow ledged . A spec ia l thank-you to M r . J u l i a n R a d l e i n , fo r h is ideas regard ing the use o f su l fonated char as a potent ia l catalyst for b iod iese l p roduc t ion . A n d last but not least, thank y o u to m y f a m i l y and f r iends fo r their encouragement and support w h i l e pursu ing this degree. i x 1 Introduction Recen t concerns over d i m i n i s h i n g foss i l fue l supp l ies and r i s i ng o i l p r ices , as w e l l as adverse env i ronmenta l and h u m a n heal th impacts f r o m the use o f pe t ro leum fue l have p rompted cons iderab le interest i n research and deve lopment o f fuels f r o m renewable resources, such as b iod iese l and e thanol . B i o d i e s e l is a very attractive al ternat ive f ue l , as it has a number o f advantages over conven t iona l d iese l fue l . It is de r i ved f r o m a renewab le , domest ic resource and can therefore reduce re l iance o n fo re ign pe t ro leum impor ts . B i o d i e s e l reduces net ca rbon d i o x i d e emiss ions by 7 8 % on a l i f e - cyc le bas is w h e n compared to conven t iona l d iese l fue l ( T y s o n 2001) . It has a lso been s h o w n to have dramat ic improvements on engine exhaust em iss ions . F o r instance, c o m b u s t i o n o f neat b i od iese l decreases ca rbon m o n o x i d e ( C O ) emiss ions by 4 6 . 7 % , part iculate matter em iss ions by 6 6 . 7 % and unburned hydrocarbons by 4 5 . 2 % (Schumacher et a l . 2001) . B i o d i e s e l can be used in a regular d iese l eng ine w i t h l i t t le to no eng ine mod i f i ca t i ons requ i red . B i o d i e s e l is safer to transport due to its h igher f lash po in t than d iese l fue l . Las t l y , b i od iese l is b iodegradab le and non - tox i c , m a k i n g it usefu l fo r t ransportat ion app l i ca t ions i n h i gh l y sens i t i ve env i ronments , such as mar ine ecosys tems and m i n i n g enc losures. H o w e v e r , b iod iese l is not w i thout its d isadvantages. These inc lude reduced energy content on per mass basis (this is due to the presence o f o x y g e n i n the fuel) w h i c h leads to l o w e r p o w e r and torque, as w e l l as h igher fue l consumpt ion . A d d i t i o n a l l y , combus t i on o f b iod iese l has been s h o w n to cause a s l ight increase in N O x f o rmat ion (Schumacher et a l . 2 0 0 1 ; D o r a d o et a l . 2003) . A s s h o w n i n E q u a t i o n 1.1, b iod iese l (def ined by the A s s o c i a t i o n for Standards and Tes t i ng o f Ma te r i a l s as m o n o - a l k y l esters o f l ong cha in fatty ac ids) is usua l l y p roduced by the t ransester i f icat ion o f a l i p i d feedstock. Transester i f i ca t ion is the revers ib le react ion o f a fat or o i l (both o f w h i c h are c o m p o s e d o f t r ig lycer ides and free fatty ac ids) w i th an a l coho l to f o r m fatty ac id a l k y l esters and g l yce ro l . S to i ch iome t r i ca l l y , the react ion requires a 3:1 a l c o h o h o i l ( A : 0 ) mo la r rat io, but because the react ion is revers ib le , excess a l coho l is added to d r i ve the e q u i l i b r i u m toward the products s ide. 1 C H 2 - O O C - R , R i - C O O - R ' C H 2 - O H I Catalyst I C H - O O C - R 2 + 3R'OH <=> R 2 - C O O - R ' + C H - O H (1.1) I I C H 2 - O O C - R 3 R 3 - C O O - R ' C H 2 - O H Glyceride Alcohol Esters Glycerol Transesterification can be alkali-, acid- or enzyme-catalyzed; however, enzyme catalysts are rarely used, as they are less effective (Ma and Hanna 1999). The reaction can also take place without the use of a catalyst under conditions in which the alcohol is in a supercritical state (Saka and Kusdiana 2001; Demirbas 2002). Biodiesel can also be produced by esterification of fatty acid molecules, as shown in Equation 1.2. This reaction can be catalyzed be either a base or an acid or without the use of a catalyst under supercritical conditions (Kusdiana and Saka 2004). Catalyst R , -COOH + R ' O H <=> R ^ C O O - R ' + H 2 0 (1.2) Fatty acid Alcohol Ester Water Currently, the high cost of biodiesel production is the major impediment to its large scale commercialization (Canakci and Van Gerpen 2001). The high cost is largely attributed to the cost of virgin vegetable oil as feedstock, which can account for up to 75% of the final product cost (Krawczyk 1996). Exploring methods to reduce the production cost of biodiesel has been the focus of much recent research. One method involves replacing a virgin oil feedstock with a waste cooking oil feedstock. The costs of waste cooking oil are estimated to be less than half of the cost of virgin vegetable oils (Canakci and Van Gerpen 2001). Furthermore, utilizing waste cooking oil has the advantage of removing a significant amount of material from the waste stream - as of 1990, it was estimated that at least 2 billion pounds of waste grease was produced annually in the United States (Canakci and Van Gerpen 2001). 1.1 Transesterification research Biodiesel related research has progressed from initial attempts to synthesize the alkyl-ester product through a simple base catalyzed reaction of pure vegetable oil to more sophisticated attempts at bringing production costs down through less expensive feedstocks, different 2 catalysts (such as homogeneous and heterogeneous ac id catalysts) and react ion cond i t ions (such as the react ion o f the l i p i d feedstock w i th a supercr i t ica l a l coho l ) . 1.1.1 Homogeneous alkali-catalyzed transesterification Transester i f i ca t ion ca ta lyzed by a base such as N a O H or K O H has been ex tens ive ly s tudied and reported (F reedman et a l . 1984; N o u r e d d i n i and Z h u 1997; M a et a l . 1998; K o m e r s et a l . 2 0 0 1 ; D o r a d o et a l . 2 0 0 2 ; D o r a d o et a l . 2004) and o p t i m u m cond i t ions at a tmospher ic pressure (60°C , 1 wt .% catalyst , 6:1 A : 0 mo la r rat io), are w e l l k n o w n (F reedman et a l . 1984). A d d i t i o n a l l y , the k ine t ics o f the react ion have been reported (F reedman et a l . 1986; N o u r e d d i n i and Z h u 1997) as f o l l o w i n g a second order react ion m e c h a n i s m , through two d is t inct react ion phases. T h e react ion rate is in i t ia l l y con t ro l l ed b y mass transfer be tween the a l coho l and o i l phases, and is then con t ro l led by k ine t ic l im i ta t ions as it approaches e q u i l i b r i u m . In order to prevent sapon i f i ca t ion (soap fo rmat ion) du r ing the react ion w h i c h leads to d i f f i cu l t y du r ing downs t ream pur i f i ca t i on , the free fatty ac id ( F F A ) and water content o f the feed must be b e l o w 0.5 wt .% and 0.05 wt.%, respect ive ly (F reedman et a l . 1984). Because o f these l im i ta t ions , on ly pure vegetable o i l feeds are appropr iate for a l ka l i - ca ta l yzed t ransester i f icat ion w i thout ex tens ive pretreatment. 7.7.2 Homogeneous acid-catalyzed transesterification A homogeneous ac id -ca ta lyzed process can be e m p l o y e d to take advantage o f cheaper feedstocks , such as waste c o o k i n g o i l and an ima l -based ta l l ow . T h e ac id -ca ta lyzed process can tolerate up to 5 wt .% F F A , but is sensi t ive to water content greater than 0.5 wt.%. T h e d isadvantage o f this method is that it is ex t remely s l o w at m i l d cond i t i ons : C a n a k c i and V a n G e r p e n (1999) , f ound that it took 48 hours to ach ieve a 9 8 % conve rs ion at 6 0 ° C at an A : 0 mo la r rat io o f 30:1 w h i c h are typ ica l cond i t ions fo r th is react ion. A t h igher temperatures and pressures (e.g. 100°C and 3.5 bar) react ion t imes can be substant ia l ly reduced (down to 8 h) to ach ieve 9 9 % conve rs ion ( G o f f et a l . 2004) . K i n e t i c studies o f the homogeneous ac id -ca ta lyzed react ion have been scarce compared to the base-cata lyzed react ion. F r e e d m a n et a l . (1986) invest igated the ac id ca ta lyzed t ransester i f icat ion o f soybean o i l w i t h bu tano l at 60°C . A t a 30:1 A : 0 mo la r rat io and 1 wt .% 3 catalyst l oad ing , the fo rward react ions were observed to be pseudo- f i rs t order w i t h the ove ra l l react ion occu r r i ng as a series o f consecu t i ve react ions. 1.1.3 Heterogeneously catalyzed transesterification A process e m p l o y i n g a heterogeneous catalyst is appea l ing because the ease o f catalyst separat ion f r o m the product stream prov ides an advantage over the t rad i t ional homogeneous processes. T o this end , s ign i f i cant effort has been expended to ident i fy and screen heterogeneous catalysts that have h igh potent ia l for b iod iese l p roduc t ion . 1.1.3.1 Solid base catalysts Severa l researchers have invest igated the t ransester i f icat ion propert ies o f so l i d base catalysts. K i m et a l . (2004) f ound that a y i e l d o f 7 8 % c o u l d be ach ieved after 2 hours us ing N a / N a O H / y - A 1 2 0 3 as a catalyst , at 6 0 ° C , 1 a tm and 6:1 A : 0 m o l a r rat io. Increased y i e l d o f 9 0 % was ach ieved by the add i t ion o f a coso lven t , n-hexane, w i th the A : 0 mo la r rat io o f 9 :1 . G r y g l e w i c z (1999) reported that after 2.5 hours at 6 0 ° C and 4.5:1 A : 0 mo la r rat io, c a l c i u m ox ide or c a l c i u m methox ide as catalyst gave b iod iese l y ie lds o f 9 0 % . H o w e v e r , no reports ex is t demonst ra t ing the ab i l i t y o f so l i d base catalysts to ester i fy F F A s present i n waste vegetable o i l and an ima l ta l low. 1.1.3.2 Solid acid catalysts D u e to their ab i l i t y to cata lyze both ester i f icat ion and t ransester i f icat ion react ions, a large number o f heterogeneous ac id catalysts i n c l u d i n g so l i d meta l ox ides and zeolytes have been screened for ac t iv i ty as s u m m a r i z e d i n Tab le 1 (Furu ta et a l . 2 0 0 4 ; L o p e z et a l . 2 0 0 5 ; J i tput t i et a l . 2006) . E x t e n s i v e w o r k has a lso gone in to deve lop ing and test ing catalysts for ester i f icat ion o f free fatty ac ids. M b a r a k a and Shanks (2005) des igned a mesoporous s i l i c a catalyst ( M C M - 4 1 ) w i t h spec ia l l y ta i lo red hyd rophob i c groups to prevent catalyst deact iva t ion by the water p roduced dur ing the ester i f ica t ion react ion. Furu ta et a l . (2004) tested their catalysts for es ter i f ica t ion act iv i ty , and reported that convers ions o f 1 0 0 % were ach ieved at a temperature o f 200°C i n the ester i f icat ion o f n-octano ic ac id w i t h methano l . T o d a et a l . (2005) recent ly deve loped an ac id catalyst by add ing su l fon i te groups to a carbon ske leton obta ined by p y r o l y z i n g ref ined sugar. Cata lys t ac t iv i ty was more than ha l f that o f the conven t iona l homogeneous ac id react ion , and 4 greater than that of other solid acid catalysts; however, the yield of the process was not mentioned. Research concerning heterogeneous catalysts is still in the catalyst screening stage. Studies regarding reaction kinetics, as well as improving reaction parameters have yet to be conducted. In addition, studies to determine the effects of free fatty acid concentration and water on the performance of the catalyst have been scarce. Table 1.1. Selected heterogeneous acid catalysts used for transesterification of triglycerides and their results. Reference Catalyst Feedstock Molar Temperature Pressure Time Conversion Type ratio (°C) (atm) (min) Achieved (%) (Furuta et Tungstated SBO* 40 •  300 1 90 al. 2004) zirconia Sulfated SBO 40 300 1 80 zirconia Sulfated tin SBO 40 300 1 68 oxide (Jitputti et Sulfated Palm 6 200 40.5 90.3 al. 2006) zirconia kernel oil Zinc oxide Palm kernel oil 6 200 40.5 86.1 Sulfated tin Palm 6 200 40.5 90.3 dioxide kernel oil K N O 3 / K L Palm 6 200 40.5 71.4 zeolyte kernel oil (Lopez et al. Amberlyst- Triacetin 6 60 1 480 79 2005) 15 Nafion Triacetin 6 60 1 480 33 NR50 Sulfated Triacetin 6 60 1 480 57 zirconia Tungstated Triacetin 6 60 1 480 10 zirconia Zeolyte HP Triacetin 6 60 1 480 <10 ETS-10(H) Triacetin 6 60 1 480 <10 *Soybean oil 1.1.4 Supercritical transesterification Supercritical transesterification is also a potential alternative to the standard homogenous catalytic routes. Transesterification using supercritical methanol has been shown to give nearly complete conversion in small amount of time (15 minutes) (Warabi et al. 2004). High temperatures, (up to 350°C) and large A : 0 rat ios (42:1) are requ i red to ach ieve the h i g h leve ls o f conve rs ion that have been reported ( K u s d i a n a and S a k a 2001) . In add i t ion to the h i g h conve rs ion and react ion rates, supercr i t ica l t ransester i f icat ion is appea l ing as it can tolerate feedstocks w i th very h igh contents o f F F A s and water, up to 36 wt .% and 30 wt.%, respect ive ly ( K u s d i a n a and S a k a 2004) . 1.2 Process modelling and economic assessment A n o t h e r impor tant too l fo r address ing the e c o n o m i c aspects o f b iod iese l is process m o d e l l i n g . Process m o d e l l i n g can be used to invest igate the effect o f process var iab les , such as p lant scale, raw mater ia l costs, u t i l i ty costs, product se l l i ng pr ices etc. o n the e c o n o m i c feas ib i l i t y o f the process. Bende r (1999) conduc ted a rev iew o f e c o n o m i c feas ib i l i t y studies f r o m di f ferent feedstocks such as beef ta l l ow and cano la seed o i l . H o w e v e r , these studies are l im i t ed to processes e m p l o y i n g an a l ka l i - ca ta l yzed react ion . M o r e recent ly , Z h a n g et a l . (2003a) deve loped a series o f H Y S Y S based process s imu la t ions to assess the techno log ica l feas ib i l i t y o f four d i f ferent b iod iese l p lant conf igura t ions - a homogeneous a l ka l i - ca ta l yzed pure vegetable o i l p rocess ; a two-s tep process to treat waste vegetable o i l ; a s ing le step homogeneous ac id -ca ta lyzed process to treat waste vegetable o i l ; and a homogeneous ac id -ca ta lyzed process us ing hexane ext rac t ion to pur i fy the b iod iese l . A l l fou r conf igura t ions were deemed techno log ica l l y feas ib le (i.e., they were capable o f p roduc ing b iod iese l to meet the A S T M spec i f i ca t ion fo r pur i ty , 99.65 wt .%) , but a subsequent e c o n o m i c analys is o f the four designs revea led that the one step ac id -ca ta lyzed process was the most e c o n o m i c a l l y attract ive process (Zhang et a l . 2003b) . Haas et a l . (2006) deve loped a process s imu la t ion m o d e l to est imate the costs o f b iod iese l p roduc t i on . T h e m o d e l was capab le o f p red ic t ing the effect on p roduc t ion cost g i v e n f luc tuat ions i n feedstock cost o r product se l l i ng pr ice . T h e m o d e l was a lso des igned to ca lcu la te the effects o n cap i ta l cost and p roduc t ion cost upon m o d i f i c a t i o n o f the process, such as changes i n feedstock type and cost , and process chemis t ry and techno logy . H o w e v e r , the m o d e l was l im i t ed to the t rad i t ional a l ka l i - ca ta l yzed p roduc t ion method . 1.3 Thesis objectives In order to determine whether the supercr i t i ca l methano l o r the heterogeneous ac id catalyst process is a p r o m i s i n g al ternat ive to the standard homogeneous cata ly t ic routes, the a i m o f Par t I 6 of this thesis is to develop a process flowsheet and simulation, conduct an economic analysis of each process based on the material and energy balance results reported by H Y S Y S , and carry out sensitivity analyses to optimize each process. Additionally, the sizing and economic calculations are incorporated into each simulation by way of the spreadsheet tool available in H Y S Y S . The material and energy flows, as well as some unit parameters are imported directly into the spreadsheet, thereby allowing the sizing and economic results to be updated automatically when any changes were made to the process flowsheet. Based on the outcome of the process simulations, it was desired to conduct more detailed catalytic studies of the heterogeneous catalyst. Therefore Part II of this thesis has investigated the synthesis and characterization of a heterogeneous catalyst, as well as testing its activity with respect to transesterification, investigating the effects reaction time, A : 0 molar ratio and catalyst loading on the outcome of the reaction, and the effects of free fatty acid content in the reaction mixture. 1.4 Thesis format The remainder of this thesis continues with two manuscripts. Chapter 2 reports the results on the design and assessment of four biodiesel production processes using H Y S Y S .Plant (submitted for publication in Bioresource Technology). Chapter 3 concentrates on the synthesis and testing of a new heterogeneous catalyst (in preparation for submission). Finally, the thesis is concluded in Chapter 4 with a general discussion of the results and recommendations for further research. References are presented at the end of each chapter. 7 1.5 References Bende r , M . (1999) . E c o n o m i c feas ib i l i t y r ev i ew fo r commun i t y - sca le fa rmer coopera t ives fo r b iod iese l . B i o resou rce T e c h n o l o g y 70(1) : 81 -87 . C a n a k c i , M . and V a n G e r p e n , J . (1999) . B i o d i e s e l p roduc t ion v i a ac id cata lys is . T ransac t ions o f the A S A E 42(5) : 1203-1210 . C a n a k c i , M . and V a n G e r p e n , J . (2001) . B i o d i e s e l p roduc t i on f r o m o i ls and fats w i th h igh free fatty ac ids . T ransac t ions o f the A S A E 44(6) : 1429-1436 . D e m i r b a s , A . (2002) . B i o d i e s e l f r o m vegetable o i l s v i a t ransester i f icat ion i n supercr i t ica l methano l . E n e r g y C o n v e r s i o n and M a n a g e m e n t 43(17) : 2349 -2356 . D o r a d o , M . P . , Ba l les te ros , E . , A r n a l , J . M . , G o m e z , J . and G i m e n e z , F . J . L . (2003). Tes t i ng waste o l i ve o i l methy l ester as a fue l i n a d iese l engine. E n e r g y & Fue ls 17(6): 1560- 1565. D o r a d o , M . P . , Ba l les te ros , E . , de A l m e i d a , J . A . , Sche l le r t , C , L o h r l e i n , H . P . and K r a u s e , R. (2002) . A n a l ka l i - ca ta l yzed t ransester i f icat ion process fo r h igh free fatty ac id waste o i l s . Transact ions o f the A S A E 45(3) : 525 -529 . D o r a d o , M . P . , Ba l les te ros , E . , M i t t e l b a c h , M . and L o p e z , F . J . (2004) . K i n e t i c parameters af fect ing the a l ka l i - ca ta l yzed t ransester i f icat ion process o f used o l i ve o i l . E n e r g y & Fue ls 18(5): 1457-1462. F r e e d m a n , B . , Bu t te r f ie ld , R. O . and P r y d e , E . H . (1986) . Transester i f i ca t ion k ine t ics o f soybean o i l . Jou rna l o f the A m e r i c a n O i l C h e m i s t s Soc ie ty 63(10) : 1375-1380 . F r e e d m a n , B . , P r y d e , E . H . and M o u n t s , T . L . (1984) . V a r i a b l e s a f fect ing the y ie lds o f fatty esters f r o m transester i f ied vegetable-Oils. Jou rna l o f the A m e r i c a n O i l C h e m i s t s Soc ie t y 61(10) : 1638-1643 . Fu ru ta , S . , M a t s u h a s h i , H . and A r a t a , K . (2004). B i o d i e s e l fue l p roduc t ion w i t h so l i d superac id cata lys is in f i x e d bed reactor under a tmospher ic pressure. Ca ta l ys i s C o m m u n i c a t i o n s 5(12) : 7 2 1 - 7 2 3 . G o f f , M . J . , Baue r , N . S . , L o p e s , S . , Sut ter l in , W . R. and Suppes , G . J . (2004) . A c i d - c a t a l y z e d a lcoho lys is o f soybean o i l . Jou rna l o f the A m e r i c a n O i l C h e m i s t s Soc ie t y 81(4): 4 1 5 - 420 . G r y g l e w i c z , S . (1999). Rapeseed o i l me thy l esters preparat ion us ing heterogeneous catalysts. B io resou rce T e c h n o l o g y 70(3) : 2 4 9 - 2 5 3 . H a a s , M . J . , M c A l o o n , A . J . , Y e e , W . C . and F o g l i a , T . A . (2006) . A process m o d e l to est imate b iod iese l p roduc t i on costs. B io resou rce T e c h n o l o g y 97(4) : 671 -678 . J i tput t i , J . , K i t i y a n a n , B . , Rangsunv ig i t , P . , Bunyak ia t , K . , A t tana tho , L . and Jenvan i tpan jaku l , P . (2006) . Transester i f i ca t ion o f c rude p a l m kerne l o i l and crude coconu t o i l by di f ferent so l i d catalysts. C h e m i c a l Eng inee r i ng Journa l 116(1): 61 -66 . K i m , H . J . , K a n g , B . S . , K i m , M . J . , Pa rk , Y . M . , K i m , D . K . , L e e , J . S . and L e e , K . Y . (2004) . Transester i f i ca t ion o f vegetable o i l to b iod iese l us ing heterogeneous base catalyst. Ca ta l ys i s T o d a y 9 3 - 9 5 : 315 -320 . K o m e r s , K . , M a c h e k , J . and S t l o u k a l , R. (2001). B i o d i e s e l f r o m rapeseed o i l , methano l and K O H 2. C o m p o s i t i o n o f so lu t ion o f K O H in methano l as react ion partner o f o i l . Eu ropean Journa l o f L i p i d Sc ience and T e c h n o l o g y 103(6): 359 -362 . K r a w c z y k , T . (1996). B i o d i e s e l . I N F O R M 7(8): 801 -822 . K u s d i a n a , D . and S a k a , S . (2001) . K i n e t i c s o f t ransester i f icat ion i n rapeseed o i l to b iod iese l fue l as treated i n supercr i t ica l methano l . F u e l 80(5) : 6 9 3 - 6 9 8 . 8 K u s d i a n a , D . and S a k a , S . (2004). E f fec ts o f water on b iod iese l fue l p roduc t ion b y supercr i t i ca l methano l treatment. B io resou rce T e c h n o l o g y 91(3) : 289 -295 . L o p e z , D . E . , G o o d w i n , J . G . , B r u c e , D . A . and Lo te ro , E . (2005) . Transester i f i ca t ion o f t r iacet in w i th methano l on so l i d ac id and base catalysts. A p p l i e d Ca ta l ys i s , A : G e n e r a l 295(2) : 97 -105 . M a , F. , C l e m e n t s , L . D . and H a n n a , M . A . (1998). T h e effects o f catalyst , free fatty ac ids , and water on t ransester i f icat ion o f beef ta l low. Transac t ions o f the A S A E 41(5) : 1261-1264. M a , F . R. and H a n n a , M . A . (1999) . B i o d i e s e l p roduc t ion : A rev iew . B io resou rce T e c h n o l o g y 70(1) : 1-15. M b a r a k a , I. K . and S h a n k s , B . H . (2005). D e s i g n o f mu l t i f unc t i ona l i zed mesoporous s i l i cas fo r ester i f icat ion o f fatty ac id . Journa l o f Ca ta l ys i s 229(2) : 365 -373 . N o u r e d d i n i , H . and Z h u , D . (1997). K i n e t i c s o f t ransester i f icat ion o f soybean o i l . Journa l o f the A m e r i c a n O i l C h e m i s t s Soc ie t y 74(11) : 1457-1463 . S a k a , S . and K u s d i a n a , D . (2001) . B i o d i e s e l fue l f r o m rapeseed o i l as prepared i n supercr i t ica l methano l . F u e l 80(2) : 2 2 5 - 2 3 1 . Schumacher , L . G . , M a r s h a l l , W . , K r a h l , J . , W e t h e r e l l , W . B . and G r a b o w s k i , M . S . (2001) . B i o d i e s e l em iss ions data f r o m series 60 ddc engines. Transact ions o f the A S A E 44(6) : 1465-1468 . T o d a , M . , T a k a g a k i , A . , O k a m u r a , M . , K o n d o , J . N . , H a y a s h i , S . , D o m e n , K . and H a r a , M . (2005) . G r e e n chemis t ry : B i o d i e s e l made w i th sugar catalyst . Na tu re 438(7065) : 178. T y s o n , K . S . B i o d i e s e l : H a n d l i n g and use gu ide l ines . h t tp : / /www.eere .energy .gov /b iomass /pd fs /b iod iese l hand l ing .pd f ( N o v e m b e r 28 , 2004) , W a r a b i , Y . , K u s d i a n a , D . and S a k a , S . (2004). Reac t i v i t y o f t r ig lycer ides and fatty ac ids o f rapeseed o i l i n supercr i t ica l a l coho ls . B io resou rce T e c h n o l o g y 91(3) : 283 -287 . Z h a n g , Y . , D u b e , M . A . , M c L e a n , D . D . and Ka tes , M . (2003a) . B i o d i e s e l p roduc t ion f r o m waste c o o k i n g o i l : 1. P rocess des ign and techno log ica l assessment. B io resou rce T e c h n o l o g y 89(1) : 1-16. Z h a n g , Y . , D u b e , M . A . , M c L e a n , D . D . and Ka tes , M . (2003b) . B i o d i e s e l p roduc t ion f r o m waste c o o k i n g o i l : 2. E c o n o m i c assessment and sens i t iv i ty ana lys is . B io resou rce T e c h n o l o g y 90(3) : 229 -240 . 9 2 Assessment of Four Continuous Biodiesel Production Processes using HYSYS-Plant1 2.1 Introduction and background Recen t concerns ove r d i m i n i s h i n g foss i l fue l supp l ies and r i s i ng o i l p r ices , as w e l l as adverse env i ronmenta l and human heal th impacts f r o m the use o f pe t ro leum fue l have p romp ted cons iderab le interest i n research and deve lopment o f fue ls f r o m renewable resources, such as b iod iese l and ethanol . B i o d i e s e l is a very attract ive al ternat ive f ue l , as it is der i ved f r o m a renewab le , domest i c resource and can therefore reduce re l iance o n fo re ign pe t ro leum impor ts . B i o d i e s e l reduces net carbon d i o x i d e emiss ions by 7 8 % on a l i f e -cyc le bas is w h e n compared to convent iona l d iese l fue l ( T y s o n 2001) . It has a lso been s h o w n to have dramat ic improvements on eng ine exhaust em iss ions . F o r instance, combus t i on o f neat b iod iese l decreases ca rbon m o n o x i d e ( C O ) em iss ions by 4 6 . 7 % , part iculate matter em iss ions by 6 6 . 7 % and unburned hydrocarbons by 4 5 . 2 % (Schumache r et a l . 2001) . A d d i t i o n a l l y , b iod iese l is b iodegradab le and non- tox i c , m a k i n g it usefu l fo r t ransportat ion app l ica t ions i n h i gh l y sensi t ive env i ronments , such as mar ine ecosystems and m i n i n g enc losures. A s s h o w n i n E q u a t i o n 2 .1 , b iod iese l (a l ky l ester) is usua l l y p roduced by the t ransester i f icat ion o f a l i p i d feedstock. Transester i f i ca t ion is the revers ib le react ion o f a fat o r o i l (both o f w h i c h are c o m p o s e d o f t r ig lycer ides and free fatty ac ids) w i th an a l coho l to f o r m fatty ac id a l k y l esters and g l yce ro l . S to i ch iome t r i ca l l y , the react ion requires a 3:1 mo la r A : 0 rat io, but because the react ion is revers ib le , excess a l coho l is added to d r i ve the e q u i l i b r i u m toward the products s ide. C H 2 - O O C - R , R i - C O O - R ' C H 2 - O H I Cata lys t I C H - O O C - R 2 + 3 R ' O H R 2 - C O O - R ' + C H - O H (2.1) I I C H 2 - O O C - R 3 R 3 - C O O - R ' C H 2 - O H G l y c e r i d e A l c o h o l Esters G l y c e r o l Transester i f i ca t ion can be a l k a l i - , a c i d - o r enzyme-ca ta l yzed ; however , e n z y m e catalysts are rarely used, as they are less ef fect ive ( M a and H a n n a 1999). T h e react ion can also take p lace w i thout the use o f a catalyst under cond i t ions i n w h i c h the a l coho l is i n a supercr i t ica l state (Saka and K u s d i a n a 2 0 0 1 ; D e m i r b a s 2002) . A version of this chapter has been submitted for publication. West, A . H . , Posarac, D. and Ellis, N . (2006) Assessment of Four Continuous Biodiesel Production Processes using HYSYS.Plant. Bioresource Technology. 10 Cur ren t l y , the h i g h cost o f b iod iese l p roduc t ion is the ma jo r imped imen t to its large scale c o m m e r c i a l i z a t i o n ( C a n a k c i and V a n G e r p e n 2001) . T h e h igh cost is la rge ly attr ibuted to the cost o f v i r g i n vegetable o i l as feedstock. E x p l o r i n g methods to reduce the p roduc t i on cost o f b iod iese l has been the focus o f m u c h recent research. O n e method i n v o l v e s rep lac ing a v i r g i n o i l feedstock w i t h a waste c o o k i n g o i l feedstock. T h e costs o f waste c o o k i n g o i l are est imated to be less than ha l f o f the cost o f v i r g i n vegetable o i l s ( C a n a k c i and V a n G e r p e n 2001) . Fur thermore , u t i l i z i ng waste c o o k i n g o i l has the advantage o f r e m o v i n g a s ign i f i can t amount o f mater ia l f r o m the waste stream - as o f 1990, it was est imated that at least 2 b i l l i o n pounds o f waste grease was p roduced annua l l y i n the U n i t e d States ( C a n a k c i and V a n G e r p e n 2001) . In the last few years, a number o f new p roduc t i on methods have emerged f r o m laboratory /bench-sca le research a i m e d at reduc ing the cost o f b iod iese l (Demi rbas 2 0 0 2 ; C a n a k c i and V a n G e r p e n 2 0 0 3 ; De l fo r t et. a l . 2003) . O n e such method uses a l coho l i n its supercr i t ica l state, and e l iminates the need for a catalyst . A d d i t i o n a l l y , the supercr i t ica l process requires on ly a short res idence t ime to reach h igh conve rs i on ( K u s d i a n a and S a k a 2004) . A n o t h e r op t ion is to use a so l i d catalyst to cata lyze the react ion (Furu ta et a l . 2004 ; Suppes et a l . 2 0 0 4 ; A b r e u et a l . 2005) . U s e o f a so l i d phase catalyst to p roduce b iod iese l w i l l s i m p l i f y downs t ream pur i f i ca t ion o f the b iod iese l . T h e catalyst can be separated by phys i ca l methods such as a hyd rocyc lone i n the case where a mu l t iphase reactor is used. A l t e rna t i ve l y , a fixed bed reactor w o u l d e l im ina te the catalyst r e m o v a l step ent i re ly . Z h a n g et a l . (2003a) deve loped a H Y S Y S based process s imu la t i on m o d e l to assess the techno log ica l feas ib i l i t y o f four b iod iese l p lant conf igura t ions - a homogeneous a lka l i - ca ta lyzed pure vegetable o i l p rocess ; a two-step process to treat waste vegetable o i l ; a s ing le step homogeneous ac id -ca ta lyzed process to treat waste vegetable o i l ; and a homogeneous a c i d - ca ta lyzed process us ing hexane ext ract ion to he lp pur i f y the b i od iese l . A l l fou r conf igura t ions were deemed techno log ica l l y feas ib le , but a subsequent e c o n o m i c analys is o f the four des igns revea led that the one step ac id -ca ta lyzed process was the most e c o n o m i c a l l y attractive process Z h a n g et a l . (2003b) . Haas et a l . (2006) deve loped a versat i le process s imu la t ion m o d e l to est imate b iod iese l p roduc t ion costs ; however , the m o d e l was l i m i t e d to a t radi t ional a l ka l i - ca ta lyzed p roduc t i on method . In order to determine whether the supercr i t i ca l methano l o r the heterogeneous ac id catalyst process is a p r o m i s i n g al ternat ive to the standard homogeneous 11 cata ly t ic routes, our a i m is to deve lop a process f lowsheet and s imu la t i on , conduc t an e c o n o m i c ana lys is o f each process based on the mater ia l and energy ba lance results reported by H Y S Y S , and carry out sens i t iv i ty analyses to op t im i ze each process. A d d i t i o n a l l y , it was des i red to automate the s i z i n g and e c o n o m i c ca lcu la t ions , whence they were incorpora ted into each s imu la t ion by way o f the spreadsheet too l ava i lab le i n H Y S Y S . T h e mater ia l and energy f l o w s , as w e l l as some uni t parameters were impor ted d i rec t ly in to the spreadsheet, thereby a l l o w i n g the s i z i n g and e c o n o m i c results to be updated au tomat ica l l y when any changes were made to the process f lowsheet . A d d i t i o n a l compar i son is made to the s imu la t ion wo rk by Z h a n g et a l . (2003a) i n order to ensure that the present s imu la t ions p rov ide comparab le resul ts. T h e homogeneous a l ka l i - ca ta l yzed system has been w e l l s tud ied, and o p t i m u m cond i t ions at 1 a tm pressure ( 6 0 ° C , 1 wt .% catalyst , 6:1 A : 0 mo la r rat io) , are k n o w n (F reedman et a l . 1984). In order to prevent sapon i f i ca t ion dur ing the react ion , the free fatty ac id ( F F A ) and water content o f the feed must be b e l o w 0.5 wt .% and 0.05 wt.%, respect ive ly (Freedman et a l . 1984). Because o f these l im i ta t ions , on ly pure vegetable o i l feeds are appropr iate for a l ka l i - ca ta l yzed t ransester i f icat ion w i thout ex tens ive pretreatment. A homogeneous ac id -ca ta lyzed process can be e m p l o y e d to take advantage o f cheaper feedstocks , such as waste c o o k i n g o i l and an ima l -based ta l low. T h e ac id -ca ta lyzed process can tolerate up to 5 wt.% F F A , but is sensi t ive to water content greater than 0.5 wt.%. T h e d isadvantage o f this me thod is that it is ex t remely s l o w at m i l d cond i t i ons : C a n a k c i and V a n G e r p e n (1999) f ound that it took 48 hours to ach ieve a 9 8 % conve rs i on at 6 0 ° C at an A : 0 rat io o f 30 :1 . A t h igher temperatures and pressures (e.g. 100°C and 3.5 bar) react ion t imes can be substant ia l ly reduced (down to 8 h) to ach ieve s im i l a r a degree o f (99%) conve rs ion ( G o f f et a l . 2004) . A process us ing a heterogeneous ac id-cata lyst is appea l ing because o f the ease o f separat ion o f a so l i d catalyst. Lo te ro et a l . (2005) reports this advantage, coup led w i t h the ab i l i ty o f the a c i d func t iona l i t y to process l o w cost , h igh free fatty ac id feedstocks, w i l l y ie ld the most e c o n o m i c a l b iod iese l p roduc t ion method . A s ou t l ined i n T a b l e 2 .1 , a number o f so l i d phase catalysts have been ident i f ied that ho l d potent ia l fo r use. Research conce rn ing heterogeneous catalysts is s t i l l i n the catalyst screening.stage. Stud ies regard ing react ion k ine t i cs , as w e l l as i m p r o v i n g react ion 12 parameters have yet to be conducted. In addition, studies to determine the effects of free fatty acid concentration and water on the performance of the catalyst have been scarce. Supercritical transesterification is also a potential alternative to the standard homogenous catalytic routes. Supercritical transesterification using methanol has been shown to give nearly complete conversion in a relatively short period (15 minutes) (Warabi et al. 2004). High temperatures, (up to 350°C) and large A : 0 ratios (42:1) are required to achieve the high levels of conversion (Kusdiana and Saka 2001). In addition to the high conversion and reaction rates, supercritical transesterification is appealing as it can tolerate feedstocks with very high contents of FFAs and water, up to 36 wt.% and 30 wt.%, respectively (Kusdiana and Saka 2004). 2.2 Process simulation To assess the technological feasibility and obtain material and energy balances for a preliminary economic analysis, complete process simulations were performed. Despite some expected differences between a process simulation and real-life operation, process simulators are commonly used to provide reliable information on process operation, owing to their vast component libraries, comprehensive thermodynamic packages and advanced computational methods. H Y S Y S Plant NetVer 3.2 was used to conduct the simulation. H Y S Y S was selected as a process simulator for both its simulation capabilities and the ability to easily incorporate sizing and economic calculations within the spreadsheet tool. The first steps in developing the process simulation were selecting the chemical components for the process, as well as a thermodynamic model. Additionally, unit operations and operating conditions, plant capacity and input conditions must all be selected and specified. The unit operations, plant capacity and input conditions for the base cases, i.e., homogeneous acid and alkali-catalyzed processes, as well as distillation column operating conditions, were selected based on the research done by Zhang et al. (2003a) to ensure that each of the four processes simulated in this work could be compared in a consistent manner. The H Y S Y S library contained information for the following components used in the simulation: methanol, glycerol, sulfuric acid, sodium hydroxide, and water. Canola oil was selected as the feedstock, and represented by triolein, as oleic acid is the major fatty acid in canola oil. Accordingly, methyl-oleate, available in the H Y S Y S component library, was taken as the product of the transesterification reaction. Where a simulation required a feedstock with some 13 amount o f free fatty ac ids , o le i c ac id , a lso ava i lab le i n the H Y S Y S l ibrary , was spec i f ied as the free fatty ac id present. C o m p o n e n t s not ava i lab le i n the H Y S Y S l ibrary were spec i f i ed us ing the " H y p o M a n a g e r " too l . C a l c i u m o x i d e , c a l c i u m sul fate, phosphor i c ac i d , s o d i u m phosphate and t r io le in were a l l spec i f ied i n this manner . Spec i f i ca t i on o f a componen t requires input o f a number o f propert ies, such as no rma l b o i l i n g po in t , densi ty , mo lecu la r we ight , as w e l l as the c r i t i ca l propert ies o f the substance. S i n c e t r io le in is a c ruc i a l componen t and is i n v o l v e d i n operat ions requ i r ing data fo r l i q u i d and vapour equ i l i b r i a , great care was taken in spec i f y ing the va lues as accurate ly as poss ib le . V a l u e s fo r densi ty , b o i l i n g po in t and c r i t i ca l temperature, pressure and v o l u m e were obta ined f r o m the A S P E N P l u s componen t l ib rary and were input as 915 k g / m 3 , 846°C , 1366°C, 4 7 0 k P a , 3.09 m 3 / k m o l , respect ive ly . A d d i t i o n a l l y , the U N D F A C structure fo r t r io le in was spec i f ied i n order to p rov i de b inary in teract ion parameter est imates. O w i n g to the presence o f po la r c o m p o u n d s such as methano l and g l yce ro l i n the process, the non - random two l i q u i d ( N R T L ) the rmodynamic /ac t i v i t y m o d e l was selected fo r use as the property package fo r the s imu la t i on . S o m e b inary in teract ion parameter coef f ic ients (such as the methano l -methy l oleate coef f ic ient ) were unava i lab le i n the s imu la t i on databanks, and were est imated us ing the U N E F A C vapou r - l i qu id e q u i l i b r i u m and U N I F A C l i q u i d - l i q u i d e q u i l i b r i u m mode ls where appropr iate. P lan t capac i ty was spec i f ied at 8000 metr ic tonnes/yr b iod iese l (Zhang et a l . 2003a) . T h i s translated to vegetable o i l feeds o f rough ly 1000 kg /h fo r each process con f igura t ion . T h e process units c o m m o n to a l l conf igura t ions i nc lude reactors, d is t i l l a t ion c o l u m n s , p u m p s and heat exchangers . T h e homogeneous a c i d - and a l ka l i - ca ta l yzed processes i nc luded l i q u i d - l i qu i d ext ract ion c o l u m n s to separate the catalyst and g l yce ro l f r o m the b iod iese l . In contrast to the base case scenar ios , a grav i ty separat ion unit was i nc l uded i n the supercr i t i ca l methano l and heterogeneous ac id catalyst processes. In spite o f the ava i lab i l i t y o f k ine t ic data fo r the a l ka l i - ca ta lyzed , homogeneous ac id -ca ta lyzed and, supercr i t i ca l processes (F reedman et a l . 1986; K u s d i a n a and S a k a 2001) , the reactors were mode led as conve rs ion reactors s ince k ine t i c i n fo rma t ion fo r the heterogeneous ac id -ca ta lyzed process was unava i lab le . T h e reactors were 14 assumed to operate con t inuous ly for a l l cases. L a b - s c a l e react ion cond i t ions and conve rs i on data were ava i lab le fo r a l l p rocesses, assumed to be appropr iate fo r large-scale p roduc t i on , and set as the operat ing cond i t ions fo r each reactor. T h e f o l l o w i n g convers ions were assumed fo r each process: 97 %, 9 7 % , 9 4 % and 9 8 % for the a l ka l i , a c i d , heterogeneous and supercr i t i ca l cases, respect ive ly (Zhang et a l . 2 0 0 3 a ; W a r a b i et a l . 2 0 0 4 ; A b r e u et a l . 2005) . T h e m o n o - and d i - g lycer ide intermediates were neglected dur ing the react ion (Zhang et a l . 2003a) . M u l t i - s t a g e d is t i l l a t ion was used to recover the excess methano l , as w e l l as i n the final pur i f i ca t ion o f b iod iese l . D i s t i l l a t i o n co lumns were spec i f ied to meet or exceed the A S T M standard fo r b iod iese l pur i ty , i.e., 99 .65 wt.%. R e f l u x rat ios fo r the heterogeneous a c i d - ca ta lyzed and supercr i t ica l cases were ca lcu la ted by de te rm in ing the m i n i m u m re f lux rat io us ing a shortcut d is t i l l a t ion c o l u m n , and then m u l t i p l y i n g by 1.5 to obta in the o p t i m u m re f lux rat io ( M c C a b e et a l . 2001) . T h e methano l recovery c o l u m n s were able to operate at ambient pressures (except i n the a lka l i - ca ta lyzed case), w h i l e v a c u u m operat ion i n the methyl -ester pu r i f i ca t ion c o l u m n s was necessary to keep the temperatures o f the d is t i l la te and bot toms streams at su i tab ly l o w leve ls , as b iod iese l and g l yce ro l are subject to degradat ion at temperatures greater than 2 5 0 ° C and 150°C, respec t ive ly ( N e w m a n 1968; G o o d r u m 2002) . 2.3 Process design F o u r con t inuous processes were s imu la ted . T w o were based on an a l ka l i - ca ta l yzed t ransester i f icat ion process us ing v i r g i n vegetable o i l (Process I), and a homogeneous a c i d - ca ta lyzed process us ing a waste c o o k i n g o i l feedstock, con ta in ing 5 wt .% free fatty ac ids (Process II). T h e th i rd con f igura t ion e m p l o y e d a heterogeneous ac id-cata lyst ( H A C ) , t in(U) ox ide , in a mu l t iphase reactor fed w i th waste vegetable o i l (Process III), w h i l e the final process used a supercr i t ica l ( S C ) methano l treatment o f waste vegetable o i l to p roduce b iod iese l (Process TV). P rocess f lowsheets are presented i n F igures 2.1 to 2.4. T h e processes a l l f o l l o w e d the same general scheme. T h e vegetable o i l was t ransester i f ied i n the first step, and then sent fo r downs t ream pur i f i ca t ion . D o w n s t r e a m pur i f i ca t ion cons is ted o f the f o l l o w i n g steps: methano l recovery by d is t i l l a t ion ; catalyst neut ra l i za t ion ; g l yce ro l separat ion and catalyst r e m o v a l ; and methyl -ester pu r i f i ca t ion by d is t i l l a t ion . T a b l e 2.2 g ives detai ls for the 15 unit operat ions i n each process. Tab les 2.3 to 2.6 present the feed and product mater ia l f l o w detai ls fo r each process. A s i l lust rated i n T a b l e 2 .1 , there are a number o f key d i f ferences between the processes. T h e f i rst d i f ference is w i th regards to the catalyst r emova l method . T h e so l i d catalyst i n Process HI is r emoved by a hyd rocyc lone before methano l recovery , whereas the l i q u i d phase catalyst i n processes I is r e m o v e d by wash ing the product stream w i t h water i n a l i q u i d - l i q u i d ex t rac t ion c o l u m n . T h e ac id catalyst i n P rocess II was r emoved as a so l i d prec ip i ta te i n separator X - 1 0 0 after neut ra l iza t ion i n reactor C R V - 1 0 1 . A s in the homogeneous ac id-cata lyst p rocess, the a lka l i -ca ta lys t had to be neut ra l i zed before it c o u l d be d i sposed of. T h e heterogeneous catalyst i n Process III requ i red no neut ra l iza t ion step; it was d iscarded as a waste product . H o w e v e r , a heterogeneous catalyst has the potent ia l advantage o f be ing recyc led . T h e second ma jo r d i f ference is i n the separat ion o f g l yce ro l f r o m the b iod iese l . In Processes I and II, g l yce ro l is r e m o v e d by wash ing the product s t ream w i t h water, and co l lec ted i n the bot toms product . In Processes III and I V , g l yce ro l is separated f r o m the b iod iese l i n a three- phase separator by grav i ty set t l ing. K r a w c z y k (1996) in i t i a l l y p roposed grav i ty separat ion to remove g l yce ro l ; however , Z h a n g et a l . (2003a) ind ica ted f r o m their s imu la t i on that sat is factory separat ion c o u l d not be ach ieved by grav i ty a lone. In the present work , grav i ty separat ion was used to separate the b iod iese l f r o m the g l yce ro l , and a sat isfactory separat ion was ach ieved . N o t e , however , that the ca lcu la t ions fo r this unit operat ion are based on parameters that have been est imated by H Y S Y S and therefore m a y not t ru ly represent a real system. A d d i t i o n a l exper imenta l data are needed to ver i f y the app l i cab i l i t y and results o f the grav i ty separator, i n order to use the uni t b l o c k w i t h con f idence . In prac t ice , a grav i ty separat ion unit has been used on a p i lo t p lant scale to separate g lycero l and b iod iese l ( C a n a k c i and V a n G e r p e n 2003) . A l l processes p roduced b iod iese l at a h igher pur i ty than requi red by the A S T M standard o f 99.65 wt.%. 2.4 Equipment sizing Process equ ipment was s i zed accord ing to p r inc ip les ou t l ined i n Tu r ton et a l . (2003) and Se ide r et a l . (2003). T h e p r i nc ipa l d imens ions o f each uni t are presented i n T a b l e 2.7. T h e equ ipment s i z i n g ca lcu la t ions were conduc ted us ing the Spreadsheet too l ava i lab le w i t h i n H Y S Y S . K e y var iab les fo r uni t s i z i n g were impor ted f r o m the f lowsheet d i rec t ly to the spreadsheet. S i z i n g 16 equat ions were encoded w i t h i n the spreadsheet. There fo re any al terat ions to the f lowsheet , such as componen t f rac t ions , componen t f lowra tes , changes to the des i red recovery i n the d is t i l l a t ions c o l u m n s , etc. are au tomat ica l l y ca lcu la ted and imp lemen ted , thus e l im ina t i ng ted ious reca lcu la t ion steps. 2.4.1 Reactor vessels Reactors were s i zed for con t inuous operat ion by d i v i d i n g the res idence t ime requi rement by the feed f lowra te fo r each process. Res idence t imes were: 4 hours , 4 hours , 3 hours and 20 minutes fo r the a l ka l i - ca ta l yzed , ac id -ca ta l yzed , heterogeneous ac id -ca ta lyzed and supercr i t i ca l processes, respect ive ly . T h e vessels, were spec i f ied to have an aspect rat io o f 3 - t o - l . 2.4.2 Columns Dis t i l l a t i on c o l u m n d iameters were s i z e d b y t w o methods. A n in i t i a l d iameter was est imated f r o m the F -Fac to r M e t h o d (Luyben 2002) . If the c o l u m n d iameter was ca lcu la ted to be greater than 0.90 m (2.95 feet) it was spec i f ied as a tray tower , and thus ca lcu la ted f r o m the f l o o d i n g ve loc i t y us ing the F a i r cor re la t ion (Se ider et a l . 2003) . C o l u m n s w i t h d iameters ca lcu la ted at less than 0.9 m were spec i f ied as a packed tower. T h e d iameter o f each packed c o l u m n was ca lcu la ted f r o m the f l o o d i n g ve loc i t y obta ined f r o m the L e v a cor re la t ion (Se ider et a l . 2003) .T ray tower height was ca lcu la ted by mu l t i p l y i ng the number o f actual stages by the tray spac ing , and then inc reas ing the result b y 2 0 % to p rov ide height fo r the condenser and reboi ler . P a c k e d tower height was ca lcu la ted by m u l t i p l y i n g the height equ iva lent o f a theoret ical p late ( H E T P ) by the number o f stages ca lcu la ted fo r the tower. H E T P was assumed to equal the c o l u m n d iameter (Se ider et a l . 2003) . A s fo r the height o f a tray tower, the packed height was increased by 2 0 % . T h e l i q u i d - l i q u i d ext ract ion c o l u m n s were s i zed acco rd ing to the wo rk o f Z h a n g et a l . (2003a) . 2.4.3 Gravity separators T h e grav i ty separators i n the heterogeneous ac id -ca ta lyzed and supercr i t ica l processes were des igned as ver t ica l process vesse ls w i t h an aspect rat io o f 2 . T h e y were s i zed to a l l o w fo r cont inuous operat ion, w i th a res idence t ime o f 1 hour . 17 2.4.4 Hydrocyclone T h e in i t ia l d imens ions o f the hyd rocyc lone (used to separate the so l i d catalyst f r o m the product stream in Process III) were ca lcu la ted by the b l o c k i n H Y S Y S . T h o s e d imens ions were then man ipu la ted s l igh t ly to obta in comp le te r e m o v a l o f the catalyst i n the hyd rocyc lone under f l ow. 2.5 Economic assessment S i n c e each process was capab le o f p roduc ing b iod iese l at the requi red pur i ty , it was o f interest to conduc t an e c o n o m i c assessment to determine process v iab i l i t y , and determine i f any one process was advantageous over the others. A s w i t h the s i z i n g ca lcu la t ions , a l l the e c o n o m i c ca lcu la t ions were pe r fo rmed w i th in the H Y S Y S spreadsheet. A d d i t i o n a l l y , the values presented fo r the e c o n o m i c ana lys is are the values obta ined after pe r fo rm ing sens i t iv i ty analyses and op t im iza t i on o f each process. T h e detai ls fo r the sens i t iv i ty analyses and op t im iza t i on studies are presented i n Sec t ion 2.6 o f this paper. A l l parameters necessary to determine mater ia l and energy costs were impor ted to the spreadsheet f r o m the f lowsheet . C o s t i n g equat ions were incorpora ted d i rec t ly in to the spreadsheet as w e l l . I nd iv idua l uni t costs were ca lcu la ted , as w e l l as f igures fo r each process i n its entirety. Incorporat ing the e c o n o m i c ca lcu la t ions into the s imu la t i on a l l o w e d fo r automat ic reca lcu la t ion o f process e c o n o m i c s shou ld any process parameters, such as componen t f lowrates o r uni t operat ing cond i t ions be changed. B y integrat ing s i z i ng and e c o n o m i c ca lcu la t ions into each s imu la t i on , the potent ia l to pe r f o rm e c o n o m i c sens i t iv i ty analyses is au tomat ica l l y bu i l t - i n to each s imu la t i on . 2.5.7 Basis of calculations E a c h process was based on a p lant capac i ty o f 8000 tonnes/year b iod iese l p roduc t ion . Opera t ing hours were set at 7 9 2 0 hours /year (assuming 330 operat ing days) . B o t h the waste and pure feedstocks were assumed free o f water and s o l i d impur i t i es , to a v o i d pre-treatment o f the feed. L o w and h igh pressure steam were used as heat ing m e d i a , w h i l e water was used for c o o l i n g . E a c h process was evaluated based on total cap i ta l investment ( T C I ) , total manufac tu r ing cost ( T M C ) , and after tax rate-of-return ( A T R O R ) . T h e assessment pe r fo rmed i n this w o r k is c lass i f i ed as a "s tudy es t imate , " w i th a range o f expec ted accuracy f r o m +30% to - 2 0 % (Tur ton et a l . 2003) . W h i l e the results o f such a study w i l l l i k e l y not ref lect the f i na l cost o f cons t ruc t ing a c h e m i c a l p lant , the techn ique is usefu l f o r p r o v i d i n g a re la t ive means to compare compe t i ng processes. 18 2.5.2 Total capital investment T a b l e 2.8 g ives a b r e a k d o w n o f the total cap i ta l investment . It a lso presents the costs for the i n d i v i d u a l uni t operat ions i n each process. B a r e m o d u l e cap i ta l costs (CBM) cons is t o f the purchase cost o f a p iece o f equ ipment , mu l t i p l i ed by the bare m o d u l e factor. Purchase costs were est imated fo r each p iece o f equ ipment based on a capac i ty equat ion presented by Tu r ton et a l . (2003) i o g l 0 c ; = ^ , + J r r 2 i o g l 0 ( A ) + / i : 3 [ i og , ( ) (A ) ] 2 (2.2) where Kt is a constant spec i f i c to the uni t type and A is the capac i ty o f the unit . B a r e m o d u l e cost was ca lcu la ted f r o m CBM ~ CPFBM (2.3) where FBM is g i ven by FBM={B,+B2FMFP) (2.4) where B] and B2 are constants spec i f i c to the unit type, and FM and FP are the mater ia l and pressure factors, respect ive ly . T h e constants, Kj and fl„ as w e l l as the pressure and mater ia l factors were obta ined f r o m Tu r ton et a l . (2003) Equa t i ons 2.2 - 2.4 were encoded w i t h i n the cos t ing spreadsheet to a l l o w fo r automat ic cost updates w h e n process parameters are changed. 2.5.3 Total manufacturing cost Di rec t manufac tu r ing expenses were ca lcu la ted based on the p r i ce and consump t i on o f each c h e m i c a l and ut i l i ty . C h e m i c a l and ut i l i ty pr ices are presented i n T a b l e 2.9 and mater ia l f l o w in fo rma t ion was obta ined f r o m H Y S Y S . Opera tor salary was est imated at $47,850/year , and it was assumed that an operator w o r k e d 49 weeks/year , and there were three 8-hour shifts per day fo r the cont inuous plant (Zhang et a l . 2003b) . T a b l e 2.10 presents a b reakdown o f the components o f the total manufac tu r ing cost as w e l l as the results fo r each process. A f t e r tax rate-of-return i s . a general c r i ter ion for e c o n o m i c per fo rmance o f a c h e m i c a l p lant. It is de f ined as the percentage o f net annua l prof i t after taxes, re la t ive to the total cap i ta l investment . N e t annual prof i t after taxes (ANNP) is ha l f the net annua l prof i t (ANP) assuming a corporate tax rate o f 5 0 % . T h e results for after tax rate o f return fo r each process are s h o w n in T a b l e 2.10. A s s h o w n i n T a b l e 2.8, the t ransester i f icat ion reactor f o rms a large part o f the cap i ta l cost , espec ia l l y fo r Processes U and I V . T h e reactor i n P rocess II was requ i red to con ta in a large mater ia l f l o w at a l ong res idence t ime. T h e presence o f su l fu r i c ac id as the catalyst requ i red a 19 stainless steel reactor, resu l t ing i n a substant ia l ly h igher reactor cost . Consequen t l y the reactor i n P rocess II was m u c h more expens ive than in a l l other processes. T h e supercr i t ica l reactor was requ i red to wi ths tand a h i g h pressure, and was const ructed f r o m stainless steel to prevent ox ida t i on and co r ros ion , hence its h igh cost . D i s t i l l a t i on c o l u m n s a lso cont r ibuted a s ign i f i can t part to the cap i ta l cost o f each process. T o w e r costs for the methyl -ester pu r i f i ca t ion tower were rough ly equa l be tween the processes, as each tower was hand l i ng approx imate ly the same mater ia l f l o w s and p r o d u c i n g b iod iese l at equa l pur i t ies. T h e methano l recovery c o l u m n s in Processes I and III were the least expens ive , as they had the smal lest mater ia l f l o w requi rements. In spite o f P rocess I V hav ing the smal lest number o f uni t operat ions, P rocess ni had the smal lest total cap i ta l investment . T h i s is due to the fact that Process I V requi red a more expens ive reactor i n order to w i ths tand the h i g h pressures and co r ros ive cond i t ions associated w i t h the supercr i t ica l state o f the a l coho l , as w e l l as the larger methano l recovery tower. T h e total cap i ta l investment fo r Process I i n the present w o r k was ca lcu la ted to be $960 thousand, less than the va lue reported by Z h a n g et a l . (2003b) o f $1 .34 m i l l i o n . T h e d i f ference l ies mos t l y i n the l o w e r costs ca lcu la ted fo r the methano l recovery c o l u m n and methyl -ester pu r i f i ca t ion c o l u m n , due to the d i f ferences in s i z i ng . Resu l ts fo r the total manufac tu r ing cost o f each process are s h o w n in Tab le 2.10. T h e di rect manufac tu r ing cost represents be tween 7 1 - 7 7 % o f the total manufac tu r ing cost i n each process. T h e largest p ropor t ion o f the di rect manufac tu r ing cost is due to the o i l feedstock - up to 5 7 % fo r Process I, and around 4 3 % fo r the other processes. P rocess in has the lowest total manufac tu r ing cost . T h i s is due to both the ab i l i t y o f the process to use l o w cost waste vegetable o i l , as w e l l as the l o w e r u t i l i ty costs o f the process resu l t ing f r o m the sma l le r mater ia l streams hand led i n the process. T h e total manufac tu r ing cost o f Process I V is s l igh t ly more than that o f Process HI, o w i n g to the large energy requi rements necessary to separate the methano l f r o m the product stream. E x c e p t for P rocess III, a l l processes had a negat ive after tax rate-of-return. Process I had the lowes t A T R O R , at - 1 4 1 % , w h i l e Processes H and I V had A T R O R s at - 4 % and - 0 . 9 % , respect ive ly . T h e A T R O R for P rocess HI was 5 4 % , ind ica t ing that the process c o u l d earn a prof i t w i thout any government subs id ies . T h e va lue for A T R O R reported b y Z h a n g et a l . (2003b) fo r P rocess I was - 8 5 % w h i c h is qui te di f ferent f r o m the va lue reported i n this work . C o m p a r i n g resul ts, the ut i l i t ies consump t i on , as w e l l as the cost o f waste d i sposa l were m u c h 20 higher in the present work, leading to a greater T M C . As well, the TCI was lower, and as its value decreases, the ATROR becomes larger in magnitude. However, our rate of return for process II (-4%) was in close agreement with the value reported for the acid-catalyzed case by al of -15%. Although the difference in magnitude between the ATROR calculated for Processes I and II is larger than that reported by Zhang et al. (2003b) the relative order of Processes I and II (i.e. that Process II has an ATROR greater than that of Process I) as presented in this work is in agreement with that of Zhang et al. (2003b). As predicted by Lotero et al. (2005), the heterogeneous acid-catalyzed process was by far the most economically attractive process. 2.6 Sensitivity analyses and optimization Sensitivity analyses were conducted to determine the effect on the process of variables that had some degree of uncertainty; and to identify any operating specifications within an individual process that could be modified to improve the process. Since the conversion data for the heterogeneous acid-catalyzed and supercritical processes were taken from bench-scale research, the economics of scale may, not be accurately reflected. Thus the effect of reduced conversion on the overall process economics was examined for each process. As shown in Figure 5, conversion in the heterogeneous acid-catalyzed process must drop to approximately 85%, while conversion in the supercritical and homogeneous acid- catalyzed processes must increase to almost 100% before there is any overlap in the ATROR. From this, it is clear that even at reduced reactor conversion, the heterogeneous process will still be advantageous over the supercritical and homogeneous acid-catalyzed processes. Sensitivity analyses were performed for all processes to determine the effect of changing methanol recovery in the methanol recovery distillation column on the ATROR. In all cases except the alkali-catalyzed case, increasing the methanol recovery caused an increase in the ATROR, due to decreased methanol consumption in all cases. Methanol acts as a cosolvent (Chiu et al. 2005) increasing the solubility of biodiesel in the glycerol phases. Therefore, reducing the amount of methanol entering the three phase separator (HAC and SC processes) reduced the amount of biodiesel lost in the glycerol stream, thereby boosting ATROR for both processes. Figure 2.6 illustrates the effect of methanol recovery on ATROR for the H A C process. Methanol recovery is limited to about 85%, as the bottoms stream temperature should not exceed 150°C. In order to increase the methanol recovery, the distillation column was 21 operated under v a c u u m cond i t i ons . T h e effect o f v a c u u m pressure (and therefore cost o f the v a c u u m system) o n the A T R O R was invest igated to determine i f the cost o f the v a c u u m sys tem was offset by the increase i n revenue that results f r o m h igher methano l recovery . A s s h o w n i n F i g u r e 2.7, the add i t ion o f the v a c u u m system resul ted i n a decrease in A T R O R . H o w e v e r , as the methano l recovery was increased under v a c u u m operat ion, the A T R O R increased, i nd ica t ing the potent ia l fo r op t im i za t i on o f the c o l u m n operat ing cond i t ions to m a x i m i z e the A T R O R . S i m i l a r analyses were conduc ted fo r the homogeneous ac id ca ta lyzed and supercr i t ica l processes, but it was f ound that v a c u u m operat ion d i d not p rov i de any e c o n o m i c benef i ts , as the methano l recovery was al ready greater than 9 9 % and the bot toms temperature was w i t h i n the a l l owab le l im i t at ambient pressure operat ion. T h e H Y S Y S op t im i ze r too l was used to vary the H A C methano l recovery i n order to m a x i m i z e the A T R O R , acco rd ing to the f o l l o w i n g constra ints: bo t toms temperature < 150°C; 1 k P a < c o l u m n pressure < 100 k P a ; and 8 5 . 0 % < methano l recovery < 9 9 . 9 % . A n o p t i m u m was found at a pressure equa l o f 4 0 k P a and methano l recovery o f 9 9 . 9 % . U p o n op t im iza t i on the bot toms temperature decreased f r o m 149.9°C to 145.5°C, methano l recovery increased f r o m 8 5 % to 9 9 . 9 % and the A T R O R increased s l igh t ly f r o m 5 3 . 7 % to 5 4 . 2 % . In add i t ion to the f i nanc ia l incent ive , i n c l u d i n g a v a c u u m system reduces methano l consump t i on and e l iminates 7 9 2 0 0 kg /y r o f methano l f r o m the waste st ream, great ly reduc ing the env i ronmenta l impac t o f the process. Las t l y , the effect o f v a c u u m operat ion i n the fatty ac id methy l -ester ( F A M E ) d is t i l l a t ion c o l u m n s was invest igated for the heterogeneous ac id -ca ta lyzed and the supercr i t ica l processes, to determine i f v a c u u m operat ion w o u l d result i n a net sav ings due to a decrease in the heat ing and c o o l i n g dut ies on the c o l u m n . C o l u m n heat ing and c o o l i n g loads d i d decrease; however , the sav ings i n ut i l i t ies cost was not enough to offset the cost o f the v a c u u m sys tem, and i n c l u s i o n o f a v a c u u m system therefore decreased the A T R O R i n bo th cases. S i n c e the upper temperature l im i t o f b i od iese l d i d not exceed at ambient opera t ion a v a c u u m sys tem was deemed unnecessary for F A M E d is t i l l a t ion i n both processes. V a c u u m operat ion for F A M E d is t i l l a t ion was needed i n the homogeneous ac id -ca ta lyzed process to keep the temperature o f the d is t i l la te b e l o w 250°C . 2.7 Conclusion F o u r con t inuous processes to p roduce b iod iese l at a rate o f 8000 tonnes/year were des igned and s imu la ted i n H Y S Y S . T h e processes were as f o l l o w s : (I) a homogeneous a l ka l i - ca ta l yzed 22 process that used pure vegetable o i l as the feedstock; (II) a homogeneous ac id -ca ta lyzed process that conver ted waste vegetable o i l as the feedstock; (HI) a heterogeneous ac id -ca ta lyzed process that used waste vegetable o i l ; and ( IV ) a supercr i t i ca l non-ca ta lyzed process, that c o n s u m e d waste vegetable o i l . F r o m a techn ica l standpoint , a l l processes are capab le o f p r o d u c i n g b iod iese l that meets A S T M spec i f i ca t ions for pur i ty . T h e supercr i t i ca l process is the s imples t and has the fewest number o f unit operat ions, but requires severe operat ing cond i t ions to ach ieve a h igh conve rs ion o f the feedstock. T h e heterogeneous ac id -ca ta lyzed process has one more unit than the supercr i t ica l process (a hyd rocyc lone to r emove the so l i d catalyst) but operates at m i l d process cond i t ions . The homogeneous processes had the greatest number o f uni t operat ions, and were more comp l i ca ted , o w i n g to the d i f f i cu l t y i n r e m o v i n g the catalyst f r o m the l i q u i d phase. A n e c o n o m i c assessment revea led that the heterogeneous ac id -ca ta lyzed process has the lowes t total cap i ta l investment , o w i n g to the re la t ive ly s m a l l s izes and carbon steel cons t ruc t ion o f mos t o f the process equ ipment . R a w mater ia ls c o n s u m e d i n the process account fo r a ma jo r por t ion o f the total manufac tu r ing cost. A c c o r d i n g l y , Processes II, HI and TV have m u c h l o w e r manufac tu r ing costs than Process I. T h e large excesses o f methano l i n Processes II and TV resul ted i n m u c h h igher ut i l i ty costs than i n process HI m a k i n g process HI the on ly process to p roduce a net prof i t . T h e after tax rate o f return fo r P rocess HI was 5 4 % , w h i l e Processes I, H and I V had rates o f return o f - 1 4 4 % , - 4 % and - 0 . 9 % , respect ive ly . Sens i t i v i t y analyses were conduc ted to ident i fy any uni t operat ions were operat ing spec i f i ca t ions c o u l d be m o d i f i e d to i m p r o v e the process. Increasing methano l recovery led to a greater A T R O R . A c c o r d i n g l y , methano l recovery was set as h igh as poss ib le (>99%) before the g l yce ro l degradat ion temperature (150°C) was exceeded in the homogeneous ac id -ca ta lyzed and supercr i t ica l processes. U s e o f the op t im i ze r func t ion ind ica ted a v a c u u m system c o u l d be ins ta l led i n the heterogeneous ac id -ca ta lyzed ( H A C ) process to increase methano l recovery and consequent ly the A T R O R , w h i l e keep ing the bot toms st ream w i t h i n the temperature l im i t . A n ana lys is o f the effect o f reac t ion conve rs i on o n A T R O R revea led that even at reduced react ion conve rs ion (i.e., between 85 -93%) , the A T R O R o f the H A C process is greater than at 1 0 0 % conve rs ion o f the homogeneous ac id and supercr i t ica l processes. 23 There fore Process in, the heterogeneous ac id -ca ta lyzed process , is c lear ly advantageous ove r the other processes, as it had the h ighest rate o f return, l owes t cap i ta l investment , and techn ica l l y , was a re la t ive ly s imp le process. Fur ther research in deve lop ing the heterogeneous ac id -ca ta lyzed process fo r b iod iese l p roduc t i on is warranted. Acknowledgments T h e authors acknow ledge the f i nanc ia l support o f the Na tu ra l Sc iences and Eng inee r i ng Research C o u n c i l o f C a n a d a . 24 Table 2.1. Catalysts and reaction parameters for heterogeneously catalyzed reactions of soybean oil at 1 atm. 2004) S n O ( A b r e u et a l . 2005) Reac t i on Parameters ^ . i . . A : 0 M o l a r _ . _ . Cata lys t type ^ Tempera ture C o n v e r s i o n T i m e W 0 3 / Z r 0 2 (Furu ta et a l . 40:1 >250 °C >90% 4 h 2004) S 0 4 / S n 0 2 (Furu ta e t a l . 40:1 300°C 6 5 % 4 h 2004) S 0 4 / Z r 0 2 (Furu ta e t a l . 40:1 300°C 8 0 % 4 h 4.15:1 60°C 94.7 3 h 25 Table 2.2. Summary of unit operating conditions for each process. Alkali- Acid- Heterogeneous Supercritical Catalyzed Catalyzed Acid-Catalyzed Process (Process I) (Process II) (Process HI) (Process IV) Transesterification Catalyst NaOH H 2 S 0 4 SnO N/A Reactor Type CSTR CSTR Multiphase CSTR Temperature (°C) 60 80 60 350 Pressure (kPa) 400 400 101.3 20x103 A : 0 Ratio 6:1 50:1 4.5:1 42:1 Residence time (hr) 4 4 3 0.333 Conversion (%) 95 97 94 98 Methanol Recovery Reflux Ratio 2 2 3.99 3.42 Number of stages 6 6 14 12 Condenser/Reboiler 20/30 101.3/111 40/50 101.3/105.3 Pressure (kPa) %Recovery 94 99.2 99.9 99.3 Distillate flowrate 113.14 1687 66.33 1239.7 (kg/h) Distillate purity(%) 100 100 99.9 99.99 Catalyst Removal N / A a N / A a hydrocyclone N/A Glycerine Separation Water Water Gravity Gravity washing washing separation separation Water flowrate 11 kg/h 46 kg/h - - Catalyst Neutralization Neutralizing agent H3PO4 C a S 0 4 N/A N/A Biodiesel Purification Reflux ratio 1.85 2 2 2 Number of stages 6 10 8 8 Condenser/reboiler 10/20 10/15 101.3/111.3 101.3/111.3 Pressure (kPa) %Recovery 99.9 98.65 99.9 99.9 Final purity 99.9 99.65 99.9 99.65 26 Table 2.3. Feed and product stream information for the alkali-catalyzed process. Feed Streams Product Streams 101 105-PVO 103 401A 401 402 501 502 Temperature (°C) 25.0 25.0 25.0 Temperature (°C) 167.8 167.5 463.9 42.8 148.6 Pressure (kPa) 101.3 101.3 101.3 Pressure (kPa) 10 10 15 20 30 Molar flow (kgmol/h) 3.61 1.19 0.25 Molar flow (kgmol/h) 0.12 3.38 0.06 0.65 1.20 Mass flow (kg/h) 115.71 1050.00 10.00 Mass flow (kg/h) 4.57 1001.8 52.77 13.79 105.12 Component mass fraction Component mass fraction Methanol 1.000 0.000 0.000 Methanol 0.6114 0.0001 0.0000 0.3432 0.0001 Triolein 0.000 1.000 0.000 Glycerol 0.0005 0.0000 0.0000 0.0002 0.9865 NaOH 0.000 0.000 1.000 Triolein 0.0000 0.0001 0.9967 0.0000 0.0014 M-oleate 0.2125 0.9997 0.0033 0.0000 0.0002 NaOH 0.0000 0.0000 0.0000 0.0000 0.0000 H3P04 0.0000 0.0000 0.0000 0.0000 0.0000 Na3P04 0.0000 0.0000 0.0000 0.0000 0.0000 Water 0.1755 0.0000 0.0000 0.6565 0.0119 Table 2.4. Feed and product stream information for the homogeneous acid-catalyzed process. Feed Streams Product Streams 101 103 105 401A 401 402 501 502 Temperature (°C) 25 25 25 Temperature (°C) 130.7 234.3 502.2 23.4 226.6 Pressure (kPa) 101.3 101.3 101.3 Pressure (kPa) 35 45 55 10 15 Molar flow (kgmol/h) 3.78 1.53 1.17 Molar flow (kgmol/h) 0.65 3.42 0.05 6.59 1.10 Mass flow (kg/h) 121.2 150.06 1030.00 Mass flow (kg/h) 20.42 1002.98 33.22 155.64 101.69 Component mass fraction Component mass fraction Methanol .1.000 0.000 0.000 Methanol 0.957 0.001 0.000 0.520 0.000 Triolein 0.000 0.000 0.950 Glycerol 0.001 0.000 0.000 0.009 0.993 H 2 S 0 4 0.000 1.000 0.000 Triolein 0.000 0.001 0.889 0.000 0.007 Oleic Acid 0.000 0.000 0.050 H2S04 0.000 0.000 0.000 0.000 0.000 M-oleate 0.007 0.998 0.111 0.003 0.000 CaO 0.000 0.000 0.000 0.000 0.000 Water 0.035 0.000 0.000 0.468 0.000 Oleic Acid 0.000 0.000 0.000 0.000 0.000 t o Table 2.5. Feed and product stream information for the heterogeneous acid-catalyzed process. Feed Streams Product Streams Methanol SnO Triolein 101 103 105 302 Glycerol Out 401 402 Temperature (°C) 25.0 25.0 25.0 Temperature (°C) 25.0 203.2 535.5 Pressure (kPa) 101.3 101.3 101.3 Pressure (kPa) 50 101.3 111.3 Molar flow (kgmol/h) 3.38 0.04 1.31 Molar flow (kgmol/h) 1.22 3.38 0.07 Mass flow (kg/h) 108.3 10.54 1050.00 Mass flow (kg/h) , 100.4 989.6 59.80 Component mass fraction Component mass fraction Methanol 1.0000 0.0000 0.0000 Methanol 0.0004 0.0000 0.0000 Triolein 0.0000 0.0000 0.9500 glycerol 0.9625 0.0001 0.0001 Tin(II) oxide 0.0000 1.0000 0.0000 triolein 0.0064 0.0000 0.9835 Oleic acid 0.0000 0.0000 0.0500 M-oleate 0.0002 0.9995 0.0165 tin(II) oxide 0.0000 0.0000 0.0000 Oleic acid 0.0000 0.0000 0.0000 water 0.0304 0.0002 0.0000 Table 2.6. Feed and product stream information for the supercritical methanol process. Feed Streams Product Streams 101 Methanol 103 Triolein 302 Glycerol Out 401 402 Temperature (°C) 25 25 Temperature (°C) 25 134.5 463.7 Pressure (kPa) 100 100 Pressure (kPa) 105.3 101.3 111.3 Molar flow (kgmol/h) 3.67 1.31 Molar flow (kgmol/h) 1.44 3.62 0.03 Mass flow (kg/h) 117.8 1050.00 Mass flow (kg/h) 110.1 1039.4 20.83 Component mass fraction Component mass fraction Methanol 1.0000 0.0000 Methanol 0.0501 0.0030 0.0000 Triolein 0.0000 0.9500 Glycerol 0.9180 0.0006 0.0000 Oleic acid 0.0000 0.0500 Triolein 0.0012 0.0000 0.9947 M-oleate 0.0033 0.9960 0.0052 Oleic acid 0.0000 0.0000 0.0000 water 0.0272 0.0003 0.0000 ro oo Table 2.7. Equipment sizes for various process units in all processes. (Dimensions are diameter x height, m). A l k a l i - A c i d - Heterogeneous Superc r i t i ca l T y p e D e s c r i p t i o n C a t a l y z e d C a t a l y z e d A c i d - C a t a l y z e d Process (Process I) (Process II) (Process HI) (Process IV) Reac to r Transester i f i ca t ion 1.8x5.4 2 .1x6.3 1.2x3.64 0 .96x2 .9 Neu t ra l i za t i on 0.36x1.1 0 .5x1.5 N / A N / A C o l u m n s M e t h a n o l R e c o v e r y 0 .46x3 0 .9x8.6 0 .31x7 .4 1x8.8 F a m e Pu r i f i ca t i on 0 .9x9.5 1x8.5 0 .9x6.6 1x6.6 W a t e r W a s h i n g 0 .8x10 1x10 N / A N / A G l y c e r o l Pu r i f i ca t i on N / A 0 .5x3.7 N / A N / A Separators G r a v i t y Separators N / A N / A 1.2x2.4 1.1x2.4 H y d r o c y c l o n e N / A N / A 0 .112x1.35 N / A 29 Table 2.8. Equipment costs, fixed capital costs and total capital investments for each process. (Units: Smillions). Alkali- Acid- Solid Acid- Supercritical Type Description Catalyzed Catalyzed Catalyzed Process Reactor Transesterification 0.292 0.680 0.075 0.639 Neutralization 0.027 0.036 0 0 Columns Methanol Recovery 0.038 0.152 0.028 0.167 Fame Purification 0.102 0.076 0.095 0.146 Washing 0.084 0.113 0 0 Glycerol Purification 0 0.028 0 0 Other Gravity Separators 0 0 0.057 0.058 Heat Exchangers 0 0.079 0.079 0.109 Pumps 0.014 0.010 0.014 0.141 Others (hydrocyclone etc) 0 0 0.015 0 Total bare module cost, C B M 0.56 1.17 0.37 1.26 Contingency fee, C C F = 0.1 8 C B M 0.10 0.22 0.07 0.23 Total module cost, C J M = C B M + C C F 0.67 1.38 0.43 1.49 Auxiliary facility cost , C A C = 0.3C B M 0.17 0.35 0.11 0.38 Fixed Capital Cost, C F c = C T M + C A C 0.83 1.73 0.54 1.87 Working capital C W c = 0.15CFc 0.13 0.26 0.08 0.28 Total capital investment C T C I = C F C + C W C 0.95 1.99 0.63 2.15 30 Table 2.9. Conditions for the economic assessment of each process. (Zhang et al. 2003b) Item Specification Price ($/tonne) Chemicals Biodiesel 600 Calcium Oxide 40 Glycerine 92 wt.% 1200 85 wt.% 750 Methanol 99.85% 180 Phosphoric Acid 340 Sodium Hydroxide 4000 Sulfuric Acid 60 Tin (II) Oxide 600 Pure canola oil 500 Waste cooking oil 200 Utilities Cooling water 400 kPa 6 °C $0.007/m3 Electricity $0.062/kWh Low pressure steama 601.3 kPa 160°C $7.78/GJ High pressure steama 4201.3 kPa 254°C $19.66/GJ * Value frozen etal. 2003) 31 Table 2.10. Total manufacturing cost and after tax rate-of-return for each process. (Units: $millions). Process I Process II Process III Process rV Direct manufacturing cost Oil feedstock 4.16 1.63 1.66 1.66 Methanol 0.16 0.30 0.16 0.17 Catalyst 0.34 0.10 0.05 0.00 Operating Labour 0.58 0.58 0.58 0.58 Supervisory Labour 0.09 0.09 0.09 0.09 LP steam 0.03 0.36 0.05 0.39 HP steam 0.26 0.25 0.28 0.33 Electricity 0.00 0.00 0.00 0.00 Cooling Water 0.01 0.02 0.01 0.02 Liquid waste disposal 0.07 0.09 0.05 0.02 Solid waste disposal 0.01 0.06 0.02 0.00 Maintenance and Repairs (M&R), 6% of C r c 0.05 0.11 0.03 0.11 Operating Supplies, 15% of M & R 0.01 0.02 0.00 0.02 Lab charges, 15% of operating labour 0.09 0.09 0.09 0.09 Patents and royalties, 3% T M C 0.22 0.15 0.12 0.14 Subtotal A D M C 6.07 3.84 3.19 3.61 Indirect manufacturing cost Overhead, packaging and storage, 60% of sum of operating labour , supervision and maintenance 0.43 0.46 0.42 0.47 Local Taxes, 1.5% of C r c 0.01 0.03 0.01 0.03 Insurance, 0.5% of C F c 0.00 0.01 0.00 0.01 Subtotal, A I M C 0.43 0.47 0.42 0.47 Depreciation 10% of C r c 0.08 0.18 0.05 0.19 General expenses Administrative costs, 25% of overhead 0.11 0.12 0.10 0.12 Distribution and selling, 10% of T M C 0.73 0.48 0.39 0.46 Research &Development, 5% of T M C 0.36 0.24 0.19 0.23 Subtotal 1.20 0.84 0.69 0.81 Total production cost 7.89 5.44 4.45 5.19 Glycerine Credit 0.62 0.60 0.57 0.60 Total Manufacturing Cost, A T E 7.28 4.83 3.88 4.59 Revenue from Biodiesel 4.75 4.76 4.70 4.92 Net annual profit -2.53 -0.08 0:82 0.33 Income taxes, A I T 50% of A N P -1.26 -0.04 0.41 0.17 Net annual after tax profit, A N N P -1.26 -0.04 0.41 0.17 After tax rate of return, I = [ A N N P - A B D ] / C T C -141.74% -10.61% 58.76% -0.90% 32  Figure 2.2. Homogeneous acid—catalyzed process flowsheet (Process II) 401A T-100 Figure 2.3. Heterogeneous acid—catalyzed process flowsheet (Process III). 1D1 Methanol 103 Triolein 302 Glycerol Out T-102 Figure 2.4. Supercritical alcohol process flowsheet (Process IV). 75 45 15 -15 cc o -45 cc 1- < -75 -105 -135 -165 80 85 100 105 90 95 Reaction Conversion (%) —•—Base Catalyzed —e—Homog. Acid Cat. A Heterog. Acid Cat. X Supercritical Figure 2.5. After-tax rate of return vs. reaction conversion for all processes. 58 57 56 g 55 O 54 CC < 53 52 51 50 0.8 0.85 0.9 0.95 Fractional Methanol Recovery 180 160 140 120 i- 3 CO 100 a> Q . 80 E o> 60 m E o 40 o m 20 0 0.75 —e— ATROR - T e m p e r a t u r e Figure 2.6. A T R O R vs. methanol recovery in the methanol recovery column, H A C process. 37 57 48 • , - —r— —r— i 100 10 20 30 40 50 60 70 80 90 100 110 Operating Pressure (kPa) —e— A T R O R - » — T e m p e r a t u r e ure 2.7. ATROR vs. operating pressure in the methanol recovery column, HAC process. 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D i s t r i bu t i on o f methano l and catalysts between b iod iese l and g l yce r in phases. A I C H E Journa l 51(4) : 1274-1278 . De l fo r t , B . , H i l l i o n , G . , L e Pennec , D . , C h o d o r g e , J . A . and B o u r n a y , L . (2003). B i o d i e s e l p roduc t ion by a con t inuous process us ing a heterogeneous catalyst. Abs t rac ts o f Papers o f the A m e r i c a n C h e m i c a l Soc ie ty 226 : U 5 3 9 - U 5 3 9 . D e m i r b a s , A . (2002) . B i o d i e s e l f r o m vegetable o i l s v i a t ransester i f icat ion i n supercr i t ica l methano l . E n e r g y C o n v e r s i o n and M a n a g e m e n t 43(17) : 2349 -2356 . F r e e d m a n , B . , Bu t te r f ie ld , R. O . and P ryde , E . H . (1986) . T ranses ter i f i ca t ion k ine t i cs o f soybean o i l . Journa l o f the A m e r i c a n O i l C h e m i s t s Soc ie ty 63(10) : 1375-1380. F r e e d m a n , B . , P r yde , E . H . and M o u n t s , T . L . (1984) . Va r i ab les a f fect ing the y ie lds o f fatty esters f r o m transester i f ied vegetab le-o i ls . Jou rna l o f the A m e r i c a n O i l C h e m i s t s Soc ie t y 61(10) : 1638-1643 . Fu ru ta , S . , M a t s u h a s h i , H . and A r a t a , K . (2004) . B i o d i e s e l fue l p roduc t i on w i t h so l i d superac id cata lys is i n f i x e d bed reactor under a tmospher ic pressure. Ca ta l ys i s C o m m u n i c a t i o n s 5(12) : 7 2 1 - 7 2 3 . G o f f , M . J . , Bauer , N . S . , L o p e s , S . , Sut te r l in , W . R. and Suppes , G . J . (2004) . A c i d - c a t a l y z e d a l coho lys i s o f soybean o i l . Jou rna l o f the A m e r i c a n O i l C h e m i s t s Soc ie ty 81(4) : 4 1 5 - 420 . G o o d r u m , J . W . (2002) . V o l a t i l i t y and b o i l i n g po in ts o f b iod iese l f r o m vegetable o i l s and ta l low. B i o m a s s & B i o e n e r g y 22(3) : 2 0 5 - 2 1 1 . H a a s , M . J . , M c A l o o n , A . J . , Y e e , W . C . and F o g l i a , T . A . (2006) . A process m o d e l to est imate b iod iese l p roduc t ion costs. B io resou rce T e c h n o l o g y 97(4) : 671 -678 . K r a w c z y k , T . (1996) . B i o d i e s e l . I N F O R M 7(8): 801 -822 . K u s d i a n a , D . and S a k a , S . (2001) . K i n e t i c s o f t ransester i f icat ion in rapeseed o i l to b iod iese l fue l as treated i n supercr i t ica l methano l . F u e l 80(5): 693 -698 . K u s d i a n a , D . and S a k a , S . (2004). E f fec ts o f water on b iod iese l fue l p roduc t ion by supercr i t i ca l methano l treatment. B io resou rce T e c h n o l o g y 91(3) : 289 -295 . Lo te ro , E . , L i u , Y . J . , L o p e z , D . E . , Suwannaka rn , K . , B r u c e , D . A . and G o o d w i n , J . G . (2005) . Synthes is o f b iod iese l v i a ac id cata lys is . Industr ia l & E n g i n e e r i n g C h e m i s t r y Resea rch 44(14) : 5 3 5 3 - 5 3 6 3 . L u y b e n , W . L . (2002) . P l a n t w i d e dynam ic s imulators i n c h e m i c a l p rocess ing and con t ro l . N e w Y o r k , M a r c e l D e k k e r . M a , F . R. and H a n n a , M . A . (1999). B i o d i e s e l p roduc t ion : A rev iew . B io resou rce T e c h n o l o g y 70(1) : 1-15. M c C a b e , W . J . , S m i t h , J . C . and Harr io t t , P . (2001). U n i t operat ions o f c h e m i c a l eng ineer ing . 6th ed . 39 N e w m a n , A . A . (1968). G l y c e r o l . C l e v e l a n d , C R C Press . S a k a , S . and K u s d i a n a , D . (2001) . B i o d i e s e l fue l f r o m rapeseed o i l as prepared i n supercr i t ica l methano l . F u e l 80(2) : 2 2 5 - 2 3 1 . Schumacher , L . G . , M a r s h a l l , W . , K r a h l , J . , W e t h e r e l l , W . B . and G r a b o w s k i , M . S . (2001) . B i o d i e s e l em iss ions data f r o m series 60 ddc eng ines. T ransac t ions o f the A S A E 44(6) : 1465-1468 . Se ider , W . D. , Seader , D . and L e w i n , D . R. (2003) . P rocess des ign p r inc ip les : Synthes is , analys is and eva lua t ion . Ch iches te r , U K , J o h n W i l e y & Sons . Suppes , G . J . , D a s a r i , M . A . , D o s k o c i l , E . J . , M a n k i d y , P . J . and G o f f , M . J . (2004) . Transester i f i ca t ion o f soybean o i l w i th zeo l i te and meta l catalysts. A p p l i e d Ca ta l ys i s a- Gene ra l 257(2) : 213 -223 . Tu r ton , R. , B a i l i e , R. C , W h i t i n g , W . B . and S h a e i w i t z , J . A . (2003) . A n a l y s i s , synthesis, and des ign o f c h e m i c a l processes. U p p e r Sadd le R i v e r , N e w Jersey, Pren t i ce H a l l . T y s o n , K . S . B i o d i e s e l : H a n d l i n g and use gu ide l ines . h t tp : / /www.eere .energy .gov /b iomass /pd fs /b iod iese l hand l ing .pd f ( N o v e m b e r 28 , 2004) , W a r a b i , Y . , K u s d i a n a , D . and S a k a , S . (2004). Reac t i v i t y o f t r ig lycer ides and fatty ac ids o f rapeseed o i l i n supercr i t ica l a l coho ls . B io resou rce T e c h n o l o g y 91(3) : 283 -287 . Z h a n g , Y . , D u b e , M . A . , M c L e a n , D . D . and K a tes , M . (2003a) . B i o d i e s e l p roduc t ion f r o m waste c o o k i n g o i l : 1. P rocess des ign and techno log ica l assessment. B io resou rce T e c h n o l o g y 89(1) : 1-16. Z h a n g , Y . , D u b e , M . A . , M c L e a n , D . D . and K a tes , M . (2003b) . B i o d i e s e l p roduc t ion f r o m waste c o o k i n g o i l : 2. E c o n o m i c assessment and sens i t iv i ty ana lys is . B io resou rce T e c h n o l o g y 90(3) : 2 2 9 - 2 4 0 . 4 0 3 Characterization and Testing of Heterogeneous Catalysts for Biodiesel Production2 3.1 Introduction and background R i s i n g crude o i l p r i ces , concerns ove r d i m i n i s h i n g foss i l fue l reserves and regard fo r env i ronmenta l qual i ty , espec ia l l y i n urban areas, have a l l cont r ibuted to the large research efforts a imed at ach iev ing e c o n o m i c a l means o f p r o d u c i n g al ternat ive fue ls der i ved f r o m renewab le resources, such as b iod iese l and b ioe thano l . B i o d i e s e l ( m o n o - a l k y l esters o f l ong cha in fatty acids) is a p r o m i s i n g al ternat ive (or extender) to conven t iona l pe t ro leum based d iese l fue l . B i o d i e s e l has a number o f advantages w h e n compared w i th regular d iese l fue l . T h e f irst and foremost is that it is der i ved f r o m a renewable domest i c resource (vegetable o i l ) , and has been s h o w n to reduce ca rbon d i o x i d e emiss ions by 7 8 % (Tyson 2001) w h e n compared to d iese l fue l o n a l i fe c y c l e bas is . C o m b u s t i o n o f b iod iese l has the potent ia l to s ign i f i can t l y l o w e r greenhouse gas ( G H G ) emiss ions . F o r examp le , a 5 % b iod iese l b lend (B5 ) inst i tu ted na t ion -w ide i n C a n a d a w o u l d reduce the amount o f CO2 enter ing the atmosphere by 2.5 M T ( H o l b e i n et a l . 2004) . B i o d i e s e l conta ins no su l fur (and therefore emits no S O x w h i c h is a precursor fo r ac id ra in) , but p rov ides greater lub r i c i t y than conven t iona l d iese l f ue l , and can therefore enhance eng ine longev i t y . L a s t l y , b iod iese l is non - tox i c and b iodegrab le m a k i n g it a more env i ronmenta l l y ben ign fue l . B i o d i e s e l is p roduced f r o m the react ion o f a vegetable o i l or an ima l fat ( w h i c h are c o m p o s e d o f c o m p l e x mix tu res o f t r ig lycer ides and free fatty ac ids depend ing on the qua l i ty o f the o i l o r ta l low) and a l o w mo lecu la r we ight a l c o h o l , such as methano l , e thanol or p ropano l . M e t h a n o l is most f requent ly used as it is the least expens ive a l c o h o l . T h e react ion between the t r ig lycer ide and the a l coho l i n the presence o f a catalyst , dep ic ted i n E q u a t i o n 3 .1 , is referred to as t ransester i f icat ion. T h e react ion produces a c o m p l e x m ix tu re o f fatty ac id methy l esters (the b iod iese l p roduct , w h i c h is dependant on the vegetable o i l type), and g l yce ro l . 2 A version of this chapter is in preparation for submission for publication. West, A . H . and Ellis, N . (2006) 41 C H 2 - O O C - R , R i - C O O - R ' C H 2 - O H I . Catalyst I C H - O O C - R 2 + 3 R ' O H <=> R 2 - C O O - R ' + C H - O H (3.1) I I C H 2 - O O C - R 3 R 3 - C O O - R ' C H 2 - O H Glyceride Alcohol Esters Glycerol Biodiesel can also be produced through the reaction of free fatty acids (FFA) and alcohol in the presence of a catalyst to produce biodiesel and water (Equation 3.2). Catalyst R i - C O O H + R ' O H <=> R i - C O O - R ' + H 2 0 (3.2) Fatty acid Alcohol Ester Water This reaction becomes significant in the case where the feedstock contains high amounts of F F A which can limit the yield of the process of the Transesterification, as water can deactivate the catalyst and lead to soap formation. The transesterification reaction can be catalyzed through a number of different methods: homoegeneous alkali (Freedman et al. 1984); homogeneous acid (Canakci and Van Gerpen 1999); supercritical alcohol with no catalyst (Saka and Kusdiana 2001) ; and via heterogeneous catalysts. The homogeneous alkali-catalyzed method is the most well known and common industrial method. It provides high yields in short times at mild process conditions, but is the most expensive of the processes (Zhang et al. 2003) , since it requires a pure vegetable oil feed (which can account for up to 7 5 % of the cost of the process (Krawczyk 1996), as the base catalyzed process is highly intolerant of water and FFA in the feedstock (Freedman et al. 1984)). Homogeneous acid-catalyzed transesterification improves on the alkali-catalyzed method as it can accommodate lower quality (and therefore less expensive) feedstocks with FFA amounts up to 5 wt.%. However, at mild conditions, the process is extremely slow, and requires up to 48 hours to achieve conversions greater than 9 5 % . It also requires a large excess of methanol. Although the process is more economical than the alkali-catalyzed method (Zhang et al. 2003) , it is still disadvantageous. Furthermore, both processes require water to separate the catalyst from the product stream and a catalyst neutralization step, increasing the waste output and necessitating a more complicated process. The supercritical process eliminates the need for a 4 2 catalyst and g ives very h i g h y ie lds i n very short t imes ( W a r a b i et a l . 2004) . H o w e v e r , these advantages are offset by the h igh cost o f the equ ipment requi red to wi ths tand such h i g h pressures (Wes t et a l . 2006) . Heterogeneous cata lys is , i n par t icu lar ac id cata lys is , presents a number o f advantages suggest ing the most e c o n o m i c a l process fo r b iod iese l p roduc t i on (Lotero et a l . 2005) . Heterogeneous catalysts can be eas i ly separated f r o m the react ion m ix tu re w i thout the use o f water, do not requi re neut ra l iza t ion and can therefore be potent ia l ly reused. In add i t ion , ac id catalysts show the potent ia l to cata lyze both ester i f icat ion and t ransester i f icat ion (Furuta et a l . 2004) react ions s imu l taneous ly , a l l o w i n g l o w e r cost feedstocks to be processed. A recent process s imu la t i on conduc ted by W e s t et a l . (2006) ind ica ted the heterogeneous process to be the most e c o n o m i c a l compared w i t h the supercr i t i ca l and t radi t ional homogeneous processes. T o this end , a number o f researchers have invest igated so l i d -ac id catalysts, such as superac ids , (Furuta et a l . 2004 ; Lo te ro et a l . 2 0 0 5 ; J i tput t i et a l . 2 0 0 6 ; K i s s et a l . 2006) and as w e l l as zeoly tes and meta l ox ides (Lo te ro et a l . 2 0 0 5 ; K i s s et a l . 2006) . A l t h o u g h the range o f temperatures, pressures and feestocks studied var ied s ign i f i can t l y , ove ra l l results were pos i t i ve , w i t h most catalysts ach iev ing > 9 0 % conve rs ion . Recen t research has a lso focused on des ign ing catalysts to e f fec t ive ly cata lyze the ester i f ica t ion o f F F A s . M b a r a k a and Shanks (2005) des igned a mesoporous s i l i c a catalyst ( M C M - 4 1 ) w i t h spec ia l l y ta i lo red hyd rophob ic groups to prevent catalyst deact iva t ion by the water p roduced dur ing the ester i f ica t ion react ion. T o d a et a l . (2005) prepared a heterogeneous ac id catalyst f r o m p y r o l i z e d sugar reacted w i t h su l fu r ic ac id and demonstrated its ab i l i t y to ester i fy free fatty ac ids , a l though they d i d not report the y i e l d o f the process. Resea rch conce rn ing heterogeneous catalysts fo r t ransester i f icat ion is s t i l l i n the catalyst screen ing stage. Stud ies regard ing react ion k inet ics are few ( L o p e z et a l . 2005) , and studies a imed at i m p r o v i n g react ion parameters have yet to be conduc ted . In add i t ion , studies to determine the effects o f f ree fatty ac id concent ra t ion and water o n the per fo rmance o f the catalyst have been scarce. B a s e d o n the pos i t i ve i nd i ca t i on that the heterogeneous process was e c o n o m i c a l , S n O was selected fo r cata ly t ic exper iments to invest igate the factors a f fect ing S n O cata lyzed t ransester i f icat ion (such as A : 0 mo la r rat io, F F A content, etc.). A n o t h e r group o f exper iments was per fo rmed w i t h an ac id catalyst de r i ved f r o m py ro l ys i s char (su l fonated char) , 4 3 to test its ability to catalyze both the transesterification and esterification reactions. Fast pyrolysis processes (heating of biomass in the absence of oxygen at rapid heating rates) generally have char yields between 10-25% by weight of the feedstock (Bridgwater et al. 1999; Dynamotiv 2006). The char can either be upgraded to activated carbon or used as an energy source, as it has a heating value comparable to lignite coal. The potential for upgrading a low- value product presents an attractive prospect, and therefore sulfonated char was investigated as a catalyst in biodiesel production. 3.2 Tin(II) oxide synthesis and testing methods 3.2.1 SnO synthesis procedure Initial attempts at synthesizing SnO followed the method of (Abreu et al. 2005). Equimolar mixtures (2 mmol) of SnCl 2 dissolved in water (20 mL) and acetylacetone were mixed under basic conditions (2 mmol NaOH in the solution) and stirred with a magnetic stirrer at 40°C for 30 minutes on a hotplate (Barnstead Thermolyne Cimarec, Fisher Scientific). The mixture was then placed in a refrigerator overnight. The precipitate was isolated via vacuum filtration (Whatman #40 filter paper), dried in a dessicator overnight and then calcined at 500°C for 24 hours in air. A second method of SnO preparation followed that of Fujita et al. (1990): an acidic solution (pH 1.1, 100 mL) of hydrocholic acid and water containing 0.02 mol/L SnCl 2 and 0.6 mol/L urea was heated at 95-97°C under reflux on a hotplate under magnetic stirring for 1 hour, at which point a dark precipitate was observed to form. The precipitate was then isolated by vacuum filtration (Whatman #40) and washed with distilled water, before being dried at room temperature in a desiccator. 3.2.2 Catalyst testing Both the prepared and commercial samples of SnO were tested as catalysts under similar conditions to the work of Abreu et al. (2005), in simple batch experiments. Reactions were carried out on a hotplate with magnetic stirring under reflux at 60°C, to determine the effect of A : 0 (methanol to canola oil respectively) molar ratio and reaction time on the conversion of the reaction. Reaction products were analyzed by GC, using a Hewlett-Packard 5890 with a flame ionization detector and a DB -5 capillary column (15 m 0.32 mm ID) (Agilent Technologies). The temperature program was as follows: Initial temperature of 45°C was held for 1 minute, and then heated at a ramp rate of 5°C/min to 300°C and held for 15 minutes. The injector and 44 detector temperatures were 2 9 0 ° C and 310°C , respect ive ly , w i t h no der i v i t i za t ion o f the samples . 3.3 Tin(II) oxide results and discussion 3.3.1 Synthesis and characterization T h e f irst p rocedure as noted i n A b r e u et a l . (2005) to synthes ize S n O d i d not resul t i n any s ign i f icant y i e ld . T h e pos t -ca lc ina t ion product was an u n k n o w n substance, d u l l grey-be ige i n co lou r , i n contrast to the sh iny b lue -b lack c o l o u r o f a c o m m e r c i a l sample o f S n O as dep ic ted i n F igures 3.1 and 3.2, respect ive ly . A subsequent rev iew o f the l i terature revealed that S n C ^ w i l l precip i tate as S n O H under bas ic cond i t ions (Fu j i ta et a l . 1990). T h u s it was l i k e l y that the ca l c ined product was some f o r m o f S n O H . N e x t , the method o f Fu j i t a et a l . (1990) was adopted to success fu l l y prepare S n O as c o n f i r m e d by c o m p a r i s o n o f x - ray d i f f rac t ion patterns o f the prepared sample w i t h the l i terature (Fu j i ta et a l . 1990), and w i t h the pattern o f a c o m m e r c i a l sample s h o w n i n F igu res 3.3 and 3.4, respect ive ly . 3.3.2 Catalytic activity In i t ia l attempts to p roduce b iod iese l by react ing cano la o i l and methano l (6:1 A : 0 mo la r rat io, 60 °C , 5 wt .% catalyst under magnet ic s t i r r ing and re f lux ) fo r 3 hours i n the presence o f the synthes ized S n O samp le p roved unsuccess fu l . Subsequent attempts he ld the catalyst l oad ing constant, and increased the A : 0 mo la r rat io (9 :1 , 15:1) and react ion t imes (12 hours , 24 hours) but no conve rs ion was observed. T h e react ion m ix tu re was ana lyzed by G C upon comp le t i ng the react ions, and showed no methy l ester peaks w h e n compared to a chromatograph f r o m a pure b iod iese l sample . Fur thermore , no not iceable react ion had occur red w h e n the c o m m e r c i a l sample o f S n O was used under react ion cond i t ions ident ica l to those descr ibed above. W i t h no other mater ia l i n the l i terature or cor respondence w i t h the authors A b r e u et a l . (2005) to support the act iv i ty o f S n O in t ransester i f icat ion react ion , further attempts to p roduce b iod iese l w i t h S n O as the catalyst ceased. 3.4 Sulfonated char synthesis and testing mthods 3.4.1 Sulfonated char synthesis procedure P y r o l y z e d h a r d w o o d char samples were obta ined f r o m R e s o u r c e T rans fo rms Internat ional L t d . (Wate r loo O N . ) , E n s y n Techno log ies Inc. (Ot tawa, O N ) and D y n a m o t i v E n e r g y Sys tems C o r p . 45 detector temperatures were 2 9 0 ° C and 3 1 0 ° C , respec t ive ly , w i t h no der i v i t i za t ion o f the samples . 3.3 Tin(II) oxide results and discussion 3.3.1 Synthesis and characterization T h e f irst procedure as noted i n A b r e u et a l . (2005) to synthes ize S n O d i d not result i n any s ign i f i cant y i e l d . T h e pos t -ca lc ina t ion product was an u n k n o w n substance, d u l l grey-be ige i n co lou r , i n contrast to the sh iny b lue -b lack c o l o u r o f a c o m m e r c i a l sample o f S n O as dep ic ted i n F igures 3.1 and 3.2, respect ive ly . A subsequent r ev i ew o f the l i terature revea led that SnCl2 w i l l prec ip i ta te as S n O H unde r .bas i c cond i t ions (Fu j i ta et a l . 1990). T h u s it was l i ke l y that the ca l c i ned product was some f o r m o f S n O H . N e x t , the method o f Fu j i t a et a l . (1990) was adopted to success fu l l y prepare S n O as c o n f i r m e d by c o m p a r i s o n o f x - ray d i f f rac t ion patterns o f the prepared sample w i t h the l i terature (Fu j i ta et a l . 1990), and w i t h the pattern o f a c o m m e r c i a l samp le s h o w n in F igu res 3.3 and 3.4, respect ive ly . 3.3.2 Catalytic activity In i t ia l attempts to p roduce b iod iese l by react ing cano la o i l and methano l (6:1 A : 0 mo la r rat io, 60 °C , 5 wt .% catalyst under magnet ic s t i r r ing and re f lux ) f o r 3 hours i n the presence o f the synthes ized S n O sample p roved unsuccess fu l . Subsequent attempts he ld the catalyst l oad ing constant , and increased the A : 0 mo la r rat io (9 :1 , 15:1) and react ion t imes (12 hours , 24 hours) but no conve rs ion was observed . T h e react ion m ix tu re was ana lyzed by G C upon comp le t i ng the react ions, and showed no methy l ester peaks w h e n compared to a chromatograph f r o m a pure b iod iese l sample . Fur thermore , no not iceab le react ion had occur red w h e n the c o m m e r c i a l sample o f S n O was used under react ion cond i t ions ident ica l to those descr ibed above. W i t h no other mater ia l i n the l i terature or cor respondence w i th the authors A b r e u et a l . (2005) to support the ac t iv i ty o f S n O i n t ransester i f icat ion react ion , further attempts to p roduce b iod iese l w i t h S n O as the catalyst ceased. . 3.4 Sulfonated char synthesis and testing mthods 3.4.1 Sulfonated char synthesis procedure P y r o l y z e d h a r d w o o d char samples were obta ined f r o m R e s o u r c e T rans fo rms Internat ional L t d . (Wate r loo O N . ) , E n s y n Techno log ies Inc. (Ot tawa, O N ) and D y n a m o t i v E n e r g y Sys tems C o r p . 45 (Vancouver BC) and sulfonated according to the method of (Toda et al. 2005). 200 mL of concentrated sulfuric acid (98%, Sigma) were added to 20 g of char in a 500 mL round bottom flask. The mixture was heated to 150°C with a heating mantle (Fisher Scientific) and monitored with a temperature controller (Omega) and corrosion-resistant Type-J thermocouple (Omega) for 24 hours. After heating, the slurry was added to cool distilled water and then vacuum filtered through #40 Whatman filter paper. The char was washed with 80°C distilled water until the wash water was neutral and free from sulfate ions. Sulfate ions were tested for by precipitation by adding several drops of a 0.66 molar barium chloride solution to the wash water. Following filtration, the char was dried in an oven at 70°C for approximately 2 hours. Samples were characterized by the following techniques: surface area was measured using nitrogen adsorption at -196°C (Micromeritics A S A P 2000) and calculated with the single-point BET method; elemental composition was determined by elemental analysis (conducted by Canadian Microanalytical Services, Delta, British Columbia); catalyst structure was analyzed via X-ray diffraction (Rigaku Multiflex X-ray diffractometer, 2 kW); surface species bonded to the catalyst surface were determine by X-ray photon spectroscopy; the total and type of acidity of the catalyst were measured by pulse n-propylamine adsorption and temperature-programmed desorption, respectively; and scanning electron microscopy was used to assess the pore size of the catalysts. Three catalyst samples were prepared from three different char samples for catalytic testing. The char samples all originated from fast pyrolysis of the following feedstocks: Catalyst 1, hardwood (RTI); Catalyst 2, hardwoods and softwoods (Ensyn); Catalyst 3, wood waste, white wood, bark and shavings (DynaMotiv). 3.4.2 Sulfonated char testing procedure The sulfonated char was tested for both transesterification and esterification activities in simple batch experiments. Reactions with canola oil were investigated to test Transesterification. Waste vegetable oil (from U B C Campus Food Outlets) was used to measure the esterification activity of the catalyst. Ethanol was used in order to achieve a higher reaction temperature (due to higher boiling point compared with methanol) and therefore faster reaction (Toda et al. 2005). Reactions were carried out on a hotplate with magnetic stirring under reflux at 76°C. The reaction mixture was analyzed by GC, as described in the Section 3.2.2. Esterification was quantified by measuring the acid number before and after the reaction. Samples were 46 cent r i fuged at 5000 r p m fo r 15 minutes to a l l o w phase separat ion. T h e o i l phase was then recovered by pipette and t i trated fo r ac id va lue us ing the M e t r o h m 794 B a s i c T i t r i no automat ic t itrator. Reac t ions were pe r fo rmed to determine the effect o f react ion t ime and A : 0 m o l a r rat io (3 hours , 9 hours and 15 hours ; 3 :1 , 6:1 9 : 1 , 12:1 and 15:1) , catalyst l oad ing (1 wt .%, 2.5 wt .% and 5 wt .%), catalyst sample (Cata lys t 1, 2 o r 3) o n the reduc t ion i n ac id number . 3.5 Sulfonated char results and discussion 3.5.1 Catalyst characterization 3.5.1.1 BET surface area T h e three catalyst samples were ana lyzed for surface area us ing n i t rogen adsorpt ion at - 196°C to determine B E T s ing le po in t sur face area as s h o w n i n T a b l e 3 .1 . E a c h sample was tested i n t r ip l icate to test fo r reproduc ib i l i t y . Table 3.1. B E T surface areas for each catalyst sample. S a m p l e 2 A r e a ( m /g) Cata lys t 1 5 . 8 4 1 0 . 3 3 Cata lys t 2 1 4 . 3 8 + 1 . 5 5 Cata lys t 3 2.74 ± 0.60 W h i l e the surface area a m o n g samples var ies somewhat , they are a l l qui te l o w as is typ ica l fo r a b u l k phase, unsuppor ted catalyst. T h e surface area fo r Cata lys t 1, w h i c h was synthes ized f r o m a h a r d w o o d der i ved samp le o f py ro lys is char is comparab le to other surface area measurements reported fo r h a r d w o o d de r i ved char ( D e l i a R o c c a et a l . 1999). T h e surface areas o f the catalyst samples were a l l greater than that reported b y T o d a et a l . (2005) . T h i s is l i k e l y due to the nature o f the char substrate, w h i c h were a l l fo rms o f w o o d b iomass . T h e structure o f the char had a h igh l y c o m p l e x ne twork o f pores, channels and otherwise f ib rous r idged surfaces (observed f r o m S E M photographs) as opposed to the p lanar structure o f the sugar -der ived char ( T o d a et a l . 2005) . 3.5.1.2 Elemental Analysis E l e m e n t a l ana lys is (presented i n T a b l e 3.2) revea led the c o m p o s i t i o n o f each catalyst sample by mass per cent, a long w i t h the co r respond ing mo lecu la r f o rmu la . 47 Table 3.2. Mass per cent composition by element and molecular formula of each catalyst sample. S a m p l e C H N O S M o l e c u l a r F o r m u l a Cata lys t 1 68 .12 2.77 0.11 28.73 2 .12 CH0.48Nn.001O0.32S0.011 Cata lys t 2 55.17 2.72 0.23 31.62 1.71 CH0.59N0.004O0.43S0.008 Cata lys t 3 70.81 2.34 0.13 20.28 1.83 CH0.39N0.001O0.22S0.009 T h e mo lecu la r f o r m u l a repor ted by T o d a et a l . (2005) fo r their catalyst was CH0.45S0.01O0 .39, w h i c h , except fo r the n i t rogen content i n the samples presented here, is very c lose , i nd ica t ing the catalysts presented here have s im i l a r compos i t i ons by mass. 3.5.1.3 X-Ray Diffraction Analysis X R D exper iments showed Cata lys t 1 was an amorphous s o l i d , s im i l a r to the catalyst repor ted b y T o d a et a l . (2005). T h e X R D spectra for Cata lys t 1 is presented i n F i g u r e 3.5. X R D spectra fo r Cata lys ts 2 and 3 a lso revea led amorphous structures. 3.5.1.4 XPS Analysis X P S exper iments were conduc ted to determine the sur face species bonded to the catalyst ca rbon substrate. A b road survey scan was conduc ted between b i n d i n g energies o f 0 e V and 1350 e V . N a r r o w scans were then conduc ted i n the S 2p reg ion , C Is reg ion and the O Is reg ion . F igu re 3.6 presents the survey scan for Cata lys t 1, w h i l e F igures 3.7 and 3.8 present the nar row scans for the S 2p and C 1 s reg ions , respect ive ly . T h e nar row O l s scan is not s h o w n . T h e peak i n F i g u r e 3.7 occurs at approx imate ly 169 e V , w h i c h cor responds to the bonded sulfate groups ( S O 4 2 ) . T h i s is i n contrast to the results repor ted by T o d a et a l . (2005) , w h o ind ica ted that S O 3 H was the bonded su l fur spec ies. There are two peaks i n F i g u r e 3.8. T h e f i rst, at 285 e V , cor responds to e lementa l ca rbon , w h i c h is the substrate o f the catalyst . T h e second , (very s m a l l peak) observed in F i g u r e 3.8 at 289 e V cor responds to c a r b o x y l i c ac id groups ( C O O H ) w h i c h is i n agreement w i t h the results o f T o d a et a l . (2005) w h o a lso reported the presence o f C O O H " groups. 3.5.1.5 n-Propylamine adsorption and temperature programmed de sorption Samp les were tested fo r total ac id i ty b y pu lse chemiso rp t i on exper iments , and then subjected to a temperature p r o g r a m m e d desorp t ion ( T P D ) to determine the type o f ac id sites. Samp les were 48 pretreated by holding the reactor temperature at 250°C for two hours in order to remove water and any adsorbed species. The sample temperature was then decreased to 120°C. After the thermal conductivity detector (TCD) readings had stabilized, the pulse experiments were conducted. The procedure was as follows. The sample loop was opened for two minutes, which allowed 1 mL of He gas containing 17.57 pmol of n-propylamine to fill the loop. At the end of the two minute period, the sample was injected into the reactor, and the outlet flow of n- propylamine measured by the TCD, logged by a multimeter (Fluke) and recorded by simple data logging software (FlukeView). After the baseline returned to an acceptable level (in all experiments the baseline was allowed to return to approximately 0.016 mV rather than 0.000 mV to reduce the length of the experiment) the sample loop was opened for two minutes and allowed to fill . The injection process was then repeated, until the peaks recorded appeared identical. In each experiment, nine pulse events were recorded. The adsorption peaks were integrated to determine the area of each one. Generally speaking, the first four of the nine peaks (Figure 3.9) of the analysis showed adsorption of the n-propylamine, while peaks 5 and beyond indicated that the catalyst sample was saturated, and therefore no n-propylamine was adsorbed. The amount of n-propylamine adsorbed during peaks 1 -4 was determined by dividing the area of each peak (from 1-4) by the average area of the saturated peaks. The total adsorption was found by adding the per cent adsorbed for peaks 1-4 and then divided by the sample mass to give a normalized value. After the pulse experiments, samples were allowed to sit for 1 hour at 120°C to remove any physisorbed species. The reactor temperature was then increased at a rate of 5°C/min to 700°C and then held for 30 minutes. The results of the pulse experiments are shown in Table 3.3 below. Table 3.3. Total acidity for each catalyst sample. Sample Total acidity (Limol/g) Catalyst 1 43.3 Catalyst 2 83.4 Catalyst 3 36.3 The total acidity of the catalysts presented here is much less than that reported by Toda et al. (2005), who achieved a total acidity of 1.4 mmol/g with the catalyst prepared from sulfuric acid. It is interesting to note however, that the acidity of the prepared catalysts is similar to the acidity 49 o f the tungstated z i r c o n i a and sul fated z i r c o n i a (54 Limol/g and 94 |amol /g , respect ive ly) tested b y ( L o p e z et a l . 2005) . T h e T P D curves fo r Cata lys ts 1 and 2 are presented i n F igu res 3.10 and 3 .11 , respect ive ly . E a c h T P D curve f o l l o w s the same pattern, and each peak occurs at approx imate ly the same temperature, i nd i ca t i ng that the types o f ac id sites on each catalyst are the same. S i n c e the peaks occur at temperatures greater than 3 0 0 ° C , the ac id i ty o f the catalysts can be attr ibuted ent i re ly to B r0ns ted ac id sites ( M i c r o m e r i t i c s 2003) . T h e T P D curves were deconvo lu ted and four sub- peaks can be observed , numbered 1 to 4 o n F igures 3.10 and 3 .11. T h e so l i d l i ne shows the T C D read ing as a func t i on o f t ime ( ind icated on each T P D curve) , w h i l e the dotted l ine shows the T C D read ing as a func t ion o f temperature. T h e t ime-ser ies cu rve has been deconvo lu ted . • A b o v e 3 0 0 ° C the n -p ropy lamine decomposes to p ropy lene and a m m o n i a . In the T P D ana lys is , the NH3 peak lags the peak for p ropy lene. O f the sma l le r peaks resu l t ing f r o m the deconvo lu t i on , the f i rst two peaks (1 and 2 , ind ica ted on F igu res 3.10 and 3.11) can be attr ibuted to p ropy lene and a m m o n i a desorp t ion , respect ive ly , f r o m a weak B r0ns ted ac id site, w h i c h cor responds to the presence o f C O O H " groups as de termined by X P S . T h e th i rd and four th peaks (numbered 3 and 4 on F igu res 3.10 and 3.11) represent desorp t ion o f p ropy lene and a m m o n i a , respect ive ly , f r o m strong B r0ns ted ac id si tes, w h i c h correlates w i t h the S 0 4 2 " groups observed i n the X P S spectra fo r each catalyst . In order to check that the deconvo lu t i on gave a reasonable resul t , the rat io o f the a m m o n i a peak area d i v i d e d b y the p ropy lene peak area can be ca lcu la ted . S i n c e p ropy lene and a m m o n i a are f o r m e d i n equ imo la r amounts f r o m the decompos i t i on o f n- p ropy lam ine , the peak areas fo r each species shou ld be equa l i f the deconvo lu t i on o f the T P D curve was done correct ly . H o w e v e r , the thermal conduc t i v i t y o f a m m o n i a is s l igh t ly greater than that o f p ropy lene (0 .0409 and 0 .0324 W / m . K , respec t i ve ly ) ; therefore the area o f the a m m o n i a peak shou ld be larger by a factor o f 1.26, i.e., the rat io o f the thermal conduc t i v i t i es o f the two species. C h e c k i n g the area rat ios fo r each set o f peaks i n F i g u r e 3.10 g ives rat ios o f 1.35 fo r peaks 1 and 2 , and 1.13 fo r peaks 3 and 4 , w h i c h are w i t h i n 7 % and 1 0 % error, respect ive ly , o f the theoret ica l va lue o f 1.26. F o r catalyst 2, the area rat ios w e r e . 1.24 fo r peaks 1 and 2 ( 2 % dev ia t ion) and 1.02 fo r peaks 3 and 4 ( 1 9 % dev ia t ion) . T h i s ind icates the results o f the deconvo lu t i on are reasonable. W i t h a re l iab le deconvo lu t i on , the re la t ive amounts o f each type o f ac id site can be determined by c o m p a r i n g the areas o f the a m m o n i a peaks . In the case o f 5 0 Cata lys t 1, the peak area fo r the strong ac id sites was greater b y approx imate ly 1 7 % , i nd i ca t i ng the total ac id i ty o f the catalyst was skewed s l igh t ly i n favour o f the strong SO42" sites. T h i s is i n contrast to the resul t reported by ( T o d a et a l . 2005) , i nd i ca t i ng that 0.7 m m o l / g o f the total 1.4 m m o l / g ac id i ty c o u l d be attr ibuted to the SO3H groups incorpora ted into the catalyst. T h e T P D curve fo r Cata lys t 2 (F igure 3.11) shows a h igher p ropor t ion o f w e a k ac id sites. T h e T P D cu rve fo r Cata lys t 3 (not shown) c o u l d not be sat is factor i ly deconvo lu ted ; i.e.the rat io o f peak areas for each pa i r o f p ropy lene and a m m o n i a peaks was never sat is factor i ly c lose enough to the theoret ica l rat io o f 1.26. T h i s c o u l d be due to e m i s s i o n o f vo la t i l e components w i t h i n the char that are not present w i t h i n the other two samples. In any case, no in fo rmat ion regard ing the d is t r ibu t ion o f act ive sites was ascerta ined fo r Cata lys t 3. 3.5.1.6 SEM Experiments V i s u a l observat ion o f the catalyst samples v i a S E M was per fo rmed to ga in ins ight in to catalyst m o r p h o l o g y and pore s ize . A l l samples were observed w i t h the same acce lerat ion vo l tage o f 20 k V . A s s h o w n i n F igu res 3.12 to 3.14, the catalyst samples have a h i gh l y i r regular , convo lu ted f ib rous surface structure, w i th l i t t le regular tex tur ing. D i sc re te pores were v i s i b l e i n some images o f Cata lys t 1, (pore d imens ions are ind ica ted i n the f igure) . H o w e v e r this was a rar i ty a m o n g the samples ana lyzed . T h e samples were a lso b r ie f l y invest igated v i a energy d ispers ive X - r a y analys is ( E D X ) to attempt to locate the act ive sites o f the catalyst by ana l yz i ng the d is t r ibu t ion o f x - rays emi t ted by su l fu r a toms upon exc i tement by the e lect ron beam. H o w e v e r , the reso lu t ion o f the E D X techn iqued was not f ine enough to p inpo in t the loca t ion o f the su l fu r e lements. Un fo r tuna te ly , the S E M and E D X exper iments y ie lded l i t t le ins ight in to nei ther the nature and loca t ion o f the catalyst act ive si tes, nor the effect o f catalyst m o r p h o l o g y on cata ly t ic act iv i ty . 3.5.2 Sulfonated char catalytic activity P r e l i m i n a r y tests w i t h the su l fonated char and cano la o i l (6:1 A : 0 rat io, 3 hours) ind ica ted s l ight t ransester i f icat ion act iv i ty . G C analys is showed ethyl-ester peaks i n the react ion m ix tu re ; however , the amount was too s m a l l to be accurate ly quant i f ied . W h e n the react ion was a l l o w e d to run fo r 24 hours at a h igher A : 0 rat io (15:1) , no v i s i b l e increase i n the amount o f b i od iese l p roduced was observed , a l though G C analys is showed the fo rmat ion o f some ethyl-esters, i nd ica t ing there is some f o r m o f resistance to t ransester i f icat ion associated w i t h the use o f the su l fonated char . 51 P r e l i m i n a r y tests w i th waste vegetable o i l co l l ec ted f r o m the U B C B i o d i e s e l P i l o t P lan t were more favourab le . T h e o i l was ana lyzed fo r ac id number before and after the react ion , and was f ound that at 12:1 A : 0 mo la r rat io and 3 hours , the ac id number decreased f r o m 8.5 m g K O H / g to 4.5 m g K O H / g . A d d i t i o n a l l y , qua l i ta t ive t ransester i f icat ion ac t iv i ty was observed upon analys is o f the react ion m ix tu re by G C . S i n c e the catalyst ind ica ted favourab le ester i f ica t ion propert ies, a set o f screen ing exper iments was conduc ted to determine the effect o f A : 0 mo la r rat io, t ime and catalyst amount on ab i l i t y o f the catalyst to reduce the F F A present i n the o i l . M o l a r rat ios invest igated were 6 :1 , 9 .5 :1 , 18 :1 , 28:1 38:1 and 4 8 : 1 , and react ion t ime was changed between 3 h, 9 h and 15 h at a f i x e d catalyst amount o f 5 wt .% based on the mass o f waste vegetable o i l . T h e range o f mo la r rat ios was selected based o n the range o f rat ios tested i n the l i terature (Furu ta et a l . 2 0 0 4 ; L o p e z et a l . 2 0 0 5 ; J i tput t i et a l . 2006) . T o determine the effect o f catalyst amount on the react ion , catalyst l oad ing was set at 1 wt .%, 2 . 5 % and 5wt .% at a f i x e d A : 0 mo la r rat io o f 28:1 and t ime o f 3 hours. F igu re 3.15 shows the effect o f reac t ion t ime at a f i x e d A : 0 mo la r rat io o n the reduct ion i n F F A . E x c e p t i n the l o w mo la r rat io cases (6:1 and 9.5:1 react ions) , inc reas ing the react ion t ime a l l o w e d fo r a greater reduc t ion i n F F A content. T h e f i na l ac id number for both the 6:1 and 9.5:1 cases stayed re la t ive ly constant , w h i c h suggests that the react ion reaches e q u i l i b r i u m fa i r l y q u i c k l y , and that increase in t ime does l i t t le to fur ther d r i ve the react ion fo rward . H o w e v e r , w h e n the A : 0 mo la r rat io is increased to 18:1 and b e y o n d , a s ign i f i cant drop i n the f i na l a c i d number can be observed , suggest ing that the increased A : 0 mo la r rat io p lays an important ro le i n d r i v i n g the e q u i l i b r i u m toward the products . A b o v e the 18:1 A : 0 mo la r rat io, there is on l y a s l ight d i f fe rence between the f i na l ac id numbers that can be attr ibuted to increased mo la r rat ios, i nd ica t ing that inc reas ing the react ion t ime p lays a greater ro le i n the conve rs ion o f the F F A . F igu re 3.16 i l lustrates the effect o f A : 0 mo la r rat io at f i x e d react ion t ime o n the reduc t ion i n F F A . A t l o w mo la r rat ios, conve rs ion o f F F A is re la t ive ly l o w . H o w e v e r , it rap id l y increases as the A : 0 mo la r rat io increases f r o m 6:1 to 18 :1 , and begins to p lateau w i t h any further increases i n A : 0 mo la r rat io. T h e error bars presented on the 15 hour t r ia l g i ve a sense o f the va r iab i l i t y 52 associated with the reaction and the quantification, and indicate that there might not be any quantifiable difference in the reduction in FFA when compared between the three reaction times, as the error bars overlap the other measurements. This has positive implications in an economic sense: since similar reaction conversion can be achieved in shorter times, this permits smaller reactor residence times, decreasing the size of the reactor; allowing for greater reactant throughput, both of which improve the economics of a production process as described by West et al. (2006). Figure 3.17 presents the effect of A : 0 molar ratio on FFA conversion, specifically for the 15 hour reaction, to give a greater sense of the variability of the experiments. The curve clearly shows that increasing the A : 0 molar ratio to the maximum ratio investigated has an impact on the reduction in FFA. However, the variability of the 18:1, 28:1 and 38:1 measurements suggests that the improvement observed by increasing the A : 0 molar ratio beyond 18:1 may be difficult to accurately quantify. Figure 3.18 illustrates the effect of catalyst amount on the conversion of FFA in the W V O . Increasing the catalyst amount in the reaction mixture increases the conversion of FFA. A greater amount of catalyst increases the number of active sites available for esterification which allows the reaction conversion to increase for a given amount of time. A similar effect on reaction conversion was observed by Kiss et al. (2006) for the esterification of oleic acid with sulfated zirconia. Figure 3.19 presents the final acid number obtained at 28:1 A : 0 molar ratio, 5 wt.% catalyst after 3 hours for each catalyst sample. While Catalysts 1 and 3 performed relatively similarly under identical reaction conditions, it is curious that under the same reaction conditions Catalyst 2 could only achieve a final acid number of 1.94, i.e., double the final acid number in the Catalyst 1 and 3 reactions, especially in consideration of the higher total acidity of Catalyst 2. It is assumed that with a higher total acidity, there are a greater number of active sites available to the reactants on the catalyst, and therefore the more acidic catalyst would show greater activity. However, it may be that the type of site is an important influence on catalyst activity. Catalyst 2 indicated a higher proportion of weak acid sites (sites with the COOH" species bonded) than did Catalyst 1, and if the weak acid sites do not catalyze the reaction as quickly or effectively (i.e. 53 the pro ton o n the C O O H " group m a y be very s l o w to d issoc ia te and attack the c a r b o x y l group o n the F F A ) as the st rong ac id sites (SO42"), this m a y account fo r the lesser ac t iv i ty observed w i t h Cata lys t 2. It was also des i red to test the catalyst fo r its ab i l i t y to ca ta lyze ester i f icat ion react ions i n a W V O w i t h a very h i g h F F A content. T h e sample o f W V O used in the cata ly t ic t r ia ls was sp i ked w i t h a sma l l amount o f o l e i c -ac i d (S igma) i n order to increase the ac id number o f the W V O to 24.5 (approx imate ly 12.25 wt .% F F A ) . T h e reac t ion was run at an a l c o h o l - t o - F F A mo la r rat io o f 160:1 ( M b a r a k a and Shanks 2005) , w h i c h translates to an A : 0 m o l a r rat io o f 7 8 : 1 . A l t h o u g h these cond i t ions are h igher than what w o u l d be used i n an indust r ia l sett ing, they were chosen to p rov ide a po in t o f c o m p a r i s o n to the h i gh l y c o m p l e x catalyst prepared by M b a r a k a and Shanks (2005). T h e catalyst l oad ing was 5 wt.%, and the reac t ion t ime was 3 hours . In three separate t r ia ls, the ac id number was reduced to an average va lue o f 2.08 ± 0 . 1 9 m g K O H / g ( rough ly 1 wt .% F F A content) . In contrast, the catalyst o f M b a r a k a and Shanks (2005) was able to reduce the amount o f F F A in a 15 wt .% pa lm i t i c ac id /soybean o i l sample to approx imate ly 2 wt.%. W h i l e the authors d i d not ind icate the m a x i m u m amount o f F F A content the catalyst c o u l d rema in act ive under, it is c lear that the catalyst prepared i n this study per fo rms comparab l y w e l l to the catalyst prepared by M b a r a k a and Shanks (2005) , but w i th the advantage o f requ i r ing a cons ide rab ly s imp le r me thod o f preparat ion. 3.6 Conclusion A study into the ef fect iveness o f tin(II) o x i d e as a catalyst fo r the t ransester i f icat ion o f vegetable o i l was conduc ted . S n O was synthes ized v i a the method o f Fu j i t a et a l . (1990) after attempts to synthes ize S n O us ing the procedure descr ibed by A b r e u et a l . (2005) fa i l ed . B o t h the synthes ized sample and a c o m m e r c i a l sample o f S n O s h o w e d no cata ly t ic act iv i ty du r ing react ions run at 6 0 ° C w i t h methano l and cano la o i l under re f lux . A second catalyst , su l fonated py ro l ys i s char , was syn thes ized based o n the technique o f T o d a et a l . (2005) et a l . Charac te r i za t ion o f the catalyst revea led an i r regular , po rous carbon f r a m e w o r k w i th C O O H " and S 0 4 2 " groups bonded to the sur face. T h e total ac id i ty o f the catalyst as revea led by pu lse n -p ropy lamine exper iments (36 - 84 Limol/g) was s im i l a r to that repor ted b y L o p e z et a l . (2005) fo r sul fated z i r con ia . Ca ta l y t i c tests w i t h cano la o i l and ethanol showed on l y qual i ta t ive t ransester i f icat ion. H o w e v e r , the catalyst was very act ive i n the ester i f icat ion o f 5 4 F F A s . Expe r imen t s ind ica ted that conve rs i on o f F F A s increased w i th increased reac t ion t ime, increased a l coho l to vegetable o i l mo la r rat io, and increased catalyst l oad ing . It was f o u n d that at a catalyst l oad ing o f 5 wt .%, reac t ion t ime o f 3 hours and A : 0 rat io o f 18:1 gave the best results. S l i gh t increases i n F F A conve rs ion were observed at mo la r rat ios b e y o n d 18:1 and react ion t imes o f 9 hours and 15 hours , but the d i f ferences were not quant i f iab le due to the var iab i l i t y assoc iated w i t h the measurements . T h e catalyst was a lso tested fo r feeds w i t h h igher F F A concentrat ions. It was f ound that the catalyst was capab le o f reduc ing the amount o f F F A s f r o m 12.25 wt .% to approx imate ly 1.04 wt .% ( w h i c h cor responds to a decrease in ac id number f r o m 24.5 to 2.04 m g K O H / g ) , w h i c h was comparab le w i t h the h igh l y c o m p l e x mesoporous s i l i ca catalyst tested by ( M b a r a k a and Shanks 2005) . Su l fona ted char shows cons iderab le potent ia l fo r use as a catalyst i n b iod iese l p roduc t i on , espec ia l l y i n a context used to reduce the free fatty ac id content o f a vegetable o i l feedstock. H o w e v e r , fo r its true potent ia l to be rea l i zed , the l im i ta t ions to t ransester i f icat ion associated w i th this catalyst must be revea led and ove rcome . Future research w i l l be d i rec ted toward this goa l . Acknowledgements T h e authors gratefu l ly acknow ledge the f i nanc ia l support o f the Na tu ra l Sc iences and Eng inee r i ng Resea rch C o u n c i l , the i n - k i n d support o f D r . K e v i n J. S m i t h , M r . I A b u fo r assistance w i t h the B E T measurements , D r . X . B . L i u fo r h is t remendous he lp w i th the n- p ropy lam ine exper iments and interpretat ion o f the resul ts, D r . A j a y K . D a l a i fo r p r o v i d i n g the E n s y n and D y n a m o t i v char samples and M r . J u l i a n R a d l e i n fo r h is ideas, input and d i scuss ion w i th regard to the use o f the py ro l ys i s char as a catalyst substrate. 55 Figure 3.1. Sample of unknown substance obtained during SnO preparation via method of Abreu et al. (2005). 7000- 6000 4 5000 ^ 4000 '55 B 3000 2000 1000 l— 20 30 40 2(9) (degrees) 50 60 Figure 3.3. X R D pattern of SnO sample prepared by method of Fujita et al. (1990). 12000H 10000- 8000- <= 6000 4 4000- 2000- I— 20 ~ i 1 1 ' r~ 30 40 50 2(0) (degrees) 60 Figure 3.4. X R D pattern of commercial SnO sample. 10 20 30 40 50 60 I— 70 - r - 80 90 26 (degrees) Figure 3.5. Catalyst 1 X R D pattern. 160000 -i 140000- 120000- 100000- 80000 - 60000 - 40000 - 20000 - 0- o o o -20000 - 1400 1200 1000 800 600 400 Binding energy (eV) 200 Figure 3.6. XPS survey scan for Catalyst 1. 3200- 3000- 2800- 2600- ^ 2400- Jj 2200- c 2000- 1800- 1600- 1400H 1200 jv'V !!' 176 174 172 170 168 166 164 Binding energy (eV) 162 160 Figure 3.7. Narrow scan in S 2p region for Catalyst 1. 80000 70000 60000 >, 50000 "55 c £ 40000 _c 30000 20000 10000 0- I . o o o c o CO O 294 292 290 288 286 284 Binding energy (eV) 282 280 Figure 3.8. Narrow scan in C 1 s region for Catalyst 1. 0.005 4 0.004 • <9 0.003 • a> w Q O 0.002 4 0.001 Peaks 5-9 Peaks 1 -4 o.ooo • L i 2000 4000 6000 8000 1 10000 Time (s) Figure 3.9. n-Propylamine pulse adsorption peaks for Catalyst 1. 0.0011 n l ' 1 ' 1 1 1 ' 1 ' 1 ' r~ 100 200 300 400 500 600 700 Temperature (C) Figure 3.10. T P D curve for Catalyst 1. Ratio of weak acid sites to strong acid sites is 0.85:1. Figure 3.12 S E M image of Catalyst 1 indicating pore sizes.  3.5- 3.0- .o> 2.5- X O E _§ 1.5- — • - 6:1 9.5:1 A 18:1 28:1 38:1 4 48:1 1.04 0.5- A 0.0- Reaction time (hours) Figure 3.15. Effect of reaction time on final acid number. Reactions were run at 5 wt.% catalyst 1 with ethanol at ArO molar ratios of 6:1, 9.5:1, 18:1, 28:1, 38:1, 48:1. 3.0- 3 hour reaction 9 hour reaction 15 hour reaction 9 2.0- 1.5- 1.0- 0.5- O 0.0- -r— 15 — i — 1 — I — 1 — | — 40 45 50 10 I 20 25 I 30 I 35 55 Alcohol to oil molar ratio Figure 3.16. Effect of A : 0 molar ratio at fixed reaction time on final acid number. 5 wt.% catalyst 1. 63 3.5- 3.0- f> 2.5 O X. °> 2.0 CD 1 1-5 3 •g 'o 1.0 < 0.5 I i 10 I 1 1 • 1 1 20 30 Alcohol to oil molar ratio 40 50 Figure 3.17. Effect of A : Q molar ratio on final acid number for the 15 hour set of reactions. 5 wt.% Catalyst 1 3.5 _ 3.0- TO I O * 2.5 TO " o n .Q 2.0 • 3 H 15- < 1.0 Catalyst amount (wt.%) Figure 3.18. Effect of catalyst amount on final acid number. 28:1 A : 0 molar ratio, ethanol, 5 wt.% catalyst 1. 64 2.50 X o O) E 2.00 1.50 to £ 1-00 "5 0.50 < 0.00 • Catalyst 1 • Catalyst 2 Catalyst 3 Catalyst sample Figure 3.19. Final acid number of reaction mixture after reaction with each catalyst sample. 3 hour reaction, 28:1 A : 0 molar ratio, 5 wt.% catalyst loading 65 3.7 References A b r e u , F . R. , A l v e s , M . B . , M a c e d o , C . C . S . , Z a r a , L . F . and Suarez , P . A . Z . (2005). N e w mul t i -phase cata ly t ic systems based on t in c o m p o u n d s act ive for vegetable o i l t ransester i f icat ion react ion . Journa l o f M o l e c u l a r Ca ta l ys i s A : C h e m i c a l 227(1-2) : 2 6 3 - 267 . B r idgwa te r , A . V . , M e i e r , D . and R a d l e i n , D . (1999) . A n o v e r v i e w o f fast py ro l ys i s o f b i omass . O r g a n i c G e o c h e m i s t r y 30(12) : 1479-1493 . C a n a k c i , M . and V a n G e r p e n , J . (1999) . B i o d i e s e l p roduc t i on v i a ac id cata lys is . T ransac t ions o f the A S A E 42(5) : 1203-1210 . D e l i a R o c c a , P . A . , C e r r e l l a , E . G . , B o n e l l i , P . R . and C u k i e r m a n , A . L . 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B io resou rce T e c h n o l o g y 91(3) : 283 -287 . W e s t , A . H . , Posa rac , D . and E l l i s , N . (2006) . Assessmen t o f four b iod iese l p roduc t ion processes us ing hysys .P lan t . B io reseou rce T e c h n o l o g y (Submi t ted fo r pub l i ca t i on February 2006) . Z h a n g , Y . , D u b e , M . A . , M c L e a n , D . D . and Ka tes , M . (2003) . B i o d i e s e l p roduc t ion f r o m waste c o o k i n g o i l : 2. E c o n o m i c assessment and sens i t iv i ty ana lys is . B i o resou rce T e c h n o l o g y 90(3) : 229 -240 . 67 4 Conclusion, General Discussion and Recommendations 4.1 General discussion Chap te r 2 featured four con t inuous processes to p roduce b iod iese l at a rate o f 8000 tonnes/year that were des igned and s imu la ted i n H Y S Y S . P l a n t , w i t h the a i m o f conduc t ing an e c o n o m i c eva lua t ion to determine w h i c h process y ie lded the most cost e f fec t ive means o f p r o d u c i n g b iod iese l . A s p rev ious l y men t ioned , the componen t t r io le in was unava i lab le i n the H Y S Y S databanks, and therefore had to be created. Ce r ta i n parameters were impor ted f r o m the A S P E N P l u s databanks where t r io le in was ava i lab le as a componen t . H o w e v e r , one key set o f parameters, the A n t o i n e ' s coef f ic ients were not ava i lab le i n A S P E N P l u s and therefore had to be est imated i n H Y S Y S . A S P E N P l u s was used to doub le check the results o f the H Y S Y S A n t o i n e ' s coe f f i c ien t es t imat ion , by us ing A S P E N P l u s to est imate its o w n set o f A n t o i n e ' s coef f i c ien ts and then g raph ing the vapour pressure as a func t i on o f temperature. D o i n g so revea led someth ing o f an anoma ly . A t l o w temperatures, it was found that the vapour pressure cu rve fo r t r io le in c rossed that o f g l yce ro l and methy l -o leate , i nd i ca t i ng it had a h igher vapour pressure, w h i c h was comp le te l y unexpected. It was expected that such a large mo lecu le w o u l d have a m u c h l o w e r vapour pressure. U s i n g the H Y S Y S A n t o i n e ' s coef f ic ients to p roduce a vapour pressure cu rve revea led the same phenomena . A rev iew o f the l i terature was under taken to obta in vapour pressure data fo r t r io le in i n order to more accurate ly pred ic t the A n t o i n e ' s coef f ic ien ts . Un fo r tuna te ly , the data were ei ther too l im i t ed or were unsat is factory and therefore unsui tab le fo r use. In l ight o f the s i tuat ion, the parameters est imated by H Y S Y S were assumed to be the best ava i lab le and used fo r the s imu la t ion . W h i l e it is des i rab le to use the most accurate cor re la t ion poss ib le s i m p l y fo r the sake o f correctness, correct parameters w i l l a lso i m p r o v e the s imu la t i on resul ts, by g i v i n g a more accurate s imu la t i on o f the methy l -o leate / t r io le in separat ion i n the second d is t i l l a t ion c o l u m n . A s s u m i n g that H Y S Y S is overp red ic t ing the vapour pressure fo r t r io le in , th is w i l l result i n h igh temperatures requ i red to separate the two componen ts , inc reas ing the energy consump t i on and therefore the cost o f the process. O f course, the oppos i te case ho lds true as w e l l . For tunate ly , this potent ia l error does not affect the re lat ive e c o n o m i c s tanding o f each process. S i n c e the mater ia l f l o w s through the re levant d is t i l la t ion c o l u m n are a l l approx imate ly equa l , the error i n terms o f energy c o n s u m p t i o n (and therefore cost) w i l l a l l be skewed to approx imate ly the same degree, l eav ing the standings unal tered. 68 Another important consideration is whether the failure to reproduce the results of Abreu et al. (2005) invalidate the conclusion that the heterogeneous process would be the most economical. To that end, a second catalyst, sulfated zirconia (S0 4 2 7Zr0 2 ) was used in the simulation. A number of researchers have confirmed the ability of S 0 4 2 7 Z r 0 2 to catalyze the transesterification of vegetable oils (Furuta et al. 2004; Lopez et al. 2005; Jitputti et al. 2006). The reaction conditions investigated by Jitputti et al. (2006) were adopted for the simulation, as they were they most rigorous in terms of temperature and pressure. The result of the simulation showed that in spite of the increased cost of the unit operations necessary for handling the large material flows and withstanding the high pressure and temperature required for the reaction, the heterogeneous process was still the most economical, although the after tax rate of return was significantly reduced from 54% in the SnO catalyzed process to 24%. This part of the work is currently in preparation for presentation at the 1 s t International Congress on Green Process Engineering in Toulouse France (2007). Based on the result from Chapter 2, Chapter 3 detailed the work that was undertaken to synthesize SnO, and then test it to assess its catalytic abilities under a variety of conditions. Unfortunately, both the commercial SnO sample obtained and the SnO sample synthesized displayed no activity during the reaction of canola oil with methanol. Discussion with Mr. J. Radlein with reference to the work of Toda et al. (2005) brought about the idea to test sulfonated pyrolysis char for transesterification acivity. Testing of the sulfonated char at an A : 0 molar ratio of 18:1 with ethanol at 76°C under reflux for 24 hours showed no visible signs of transesterification, but analysis of the reaction mixture via GC indicated the presence of some ethyl-ester. The chromatogram (not shown) also exhibited peaks associated with glycerol, di- glycerides and mono-glycerides, which would not be present if the ethyl-ester was being produced exclusively through esterification of any free fatty acids present in the oil. Due to time constraints the limitations to transesterification could not be explored. There are a number of possibilities that could explain the lack of transesterification acitivity. The first is that the catalyst may not have been acidic enough. However, the total acidity measured for the sulfonated char was comparable to the acidity of sulfated zirconia reported by Lopez et al. (2005), and the activity of sulfated zirconia is well confirmed. Catalytic studies have also shown that internal resistance to mass transfer and stearic hindrance can also limit catalyst activity when microporous catalysts are used, such as Zeolite HP (pore size 5.5 A X 5.5 A), H-ZSM5 69 and Y ( L o p e z et a l . 2 0 0 5 ; K i s s et a l . 2006) and that i n such cases any ac t iv i ty was the resul t o f surface sites. S E M exper iments were conduc ted to assess the surface character is t ics o f the catalyst , but no regu lar ly occu r r i ng pore structures c o u l d be observed . W h e r e they were f o u n d (F igure 3.12) the surface pore s ize was qui te large (>1.6 p m ) w h i c h w o u l d not present any resistance to d i f f us ion . H o w e v e r , some f o r m o f mass transfer res istance may be l i m i t i n g act iv i ty , i f perhaps the act ive sites were a l l w i t h i n the l ong f ib rous channels observed i n F i g u r e 3 .13. T o assess whether the act ive sites were located o n the surface o f the catalyst or w i t h i n the pores /channe ls , an E D X scan was per fo rmed to locate the S-con ta in ing sites. Un fo r tuna te ly the reso lu t ion was not f ine enough to p inpo in t the loca t ion o f the su l fur and no in fo rmat ion regard ing the loca t ion o f the ac t ive sites c o u l d be ga ined. A n o t h e r poss ib i l i t y m a y have been that the react ion temperature was too l o w fo r t ransester i f icat ion to occur . O the r studies (Furu ta et a l . 2 0 0 4 ; Suppes et a l . 2 0 0 4 ; J i tput t i et a l . ! I 2006) have used m u c h h igher temperatures ( > 1 5 0 ° C ) , than c o u l d be ach ieved w i th the s imp le hotplate set-up e m p l o y e d fo r this work . None the less , i f temperature is the l i m i t i n g factor , it is expected that some t ransester i f icat ion w o u l d be observed at l o w e r temperatures ( L o p e z et a l . 2005) . 4.2 Conclusions U s i n g the H Y S Y S s imula tor , process f lowsheets and energy and mater ia l ba lances were deve loped to m o d e l the processes. T h e integrated spreadsheet too l i n H Y S Y S was used to conduc t uni t operat ion s i z i n g , as w e l l as automate the e c o n o m i c ca lcu la t ions , w h i c h i nc luded a l l equ ipment cos t ing , total cap i ta l investment , total manufac tu r ing cost and after tax rate o f return. T h e processes were as f o l l o w s : (I) a homogeneous a l ka l i - ca ta l yzed process that used pure vegetable o i l as the feedstock; (II) a homogeneous ac id -ca ta lyzed process that conver ted waste vegetable o i l as the feedstock; (III) a heterogeneous ac id -ca ta lyzed process that used waste vegetable o i l ; and ( IV ) a supercr i t ica l non-ca ta lyzed process , that c o n s u m e d waste vegetable o i l . T h e supercr i t i ca l process was the s imples t and had the fewest number o f uni t operat ions, w h i l e the homogeneous processes had the greatest number o f uni t operat ions, and were the mos t comp l i ca ted , o w i n g to the d i f f i cu l t y i n r e m o v i n g the catalyst f r o m the l i q u i d phase. A n e c o n o m i c assessment revea led that the heterogeneous ac id -ca ta lyzed process had the lowes t total cap i ta l investment and total manufac tu r ing cost. It was f ound that raw mater ia ls c o n s u m e d and the s ize o f mater ia l f l o w s , s t rongly af fected process e c o n o m i c s . A c c o r d i n g l y , Processes II, 70 I l l and I V had m u c h l o w e r manufac tu r ing costs than Process I. T h e after tax rate o f return for process III was 5 4 % , w h i l e processes I, II and I V had rates o f return o f - 1 4 4 % , - 4 % and - 0 . 9 % , respect ive ly . Sens i t i v i t y analyses were conduc ted to ident i fy any uni t operat ions where operat ing spec i f i ca t ions c o u l d be m o d i f i e d to i m p r o v e the process. It was found that i nc reas ing methano l recovery led to a greater A T R O R . A c c o r d i n g l y , me thano l recovery was set as h i g h as poss ib le (>99%) before the g l yce ro l degradat ion temperature (150°C) was exceeded in the homogeneous ac id -ca ta lyzed and supercr i t i ca l processes. U s e o f the op t im i ze r f unc t i on ind ica ted a v a c u u m system c o u l d be ins ta l led i n the H A C process to increase methano l recovery and consequent ly the A T R O R , w h i l e keeping 1 the bot toms stream w i th in the temperature l im i t . A n analys is o f the effect o f react ion conve rs ion on A T R O R revea led that even at reduced react ion conve rs ion (i.e., be tween 85 -93%) , the A T R O R o f the H A C process is greater than at 1 0 0 % conve rs ion o f the homogeneous ac id and supercr i t i ca l processes. There fo re P rocess HI , the heterogeneous a c i d - ca ta lyzed process, is c lear ly advantageous ove r the other processes, as it had the h ighest rate o f return, lowest cap i ta l investment , and techn ica l l y , was a re la t i ve ly s imp le process. Fur ther research i n deve lop ing the heterogeneous ac id -ca ta lyzed process fo r b iod iese l p roduc t ion is warranted. B a s e d on the results f r o m the H Y S Y S s imu la t i on , a study into the ef fect iveness o f tin(II) o x i d e as a catalyst fo r the t ransester i f icat ion o f vegetable o i l was conduc ted . S n O was synthes ized v i a the method o f Fu j i t a et a l . (1990) after attempts to synthes ize S n O us ing the procedure descr ibed by A b r e u et a l . (2005) had fa i led . B o t h the synthes ized samp le and a c o m m e r c i a l sample o f S n O showed no cata ly t ic ac t iv i ty du r ing react ions run at 6 0 ° C w i th methano l and cano la o i l under re f lux . A second catalyst , su l fonated py ro l ys i s char , was synthes ized based on the technique o f T o d a et a l . (2005) et a l . Charac te r i za t ion o f the catalyst revea led an i r regular , porous carbon f r amework w i t h C O O H " and SO42" groups bonded to the surface. Ca ta l y t i c tests w i th cano la o i l and ethanol showed on l y qual i ta t ive (i.e., an amount too s m a l l to be phys i ca l l y measured) t ransester i f icat ion. H o w e v e r , the catalyst was very act ive i n the es ter i f i ca t ion o f F F A s . Expe r imen t s s h o w e d that conve rs i on o f F F A s increased w i t h inc reas ing react ion t ime, i nc reas ing a l coho l to vegetable o i l 71 mo la r rat io, and inc reas ing catalyst l oad ing . It was f ound that at a catalyst l oad ing o f 5 wt .%, react ion t ime o f 3 hours and A : 0 rat io o f 18:1 gave the best resul ts. S l i gh t increases i n F F A conve rs i on were observed at mo la r rat ios b e y o n d 18:1 and react ion t imes o f 9 hours and 15 hours , but the d i f ferences were not quant i f iab le due to the var iab i l i t y assoc iated w i t h the measurements. A t h igher F F A concent ra t ions, it was f ound that the catalyst was capab le o f reduc ing the amount o f F F A s f r o m 12.25 wt .% to approx imate ly 1.04 wt.%. Su l fona ted char shows cons iderab le potent ia l fo r use as a catalyst i n b i od iese l p roduc t i on , espec ia l l y i n a context used to reduce the free fatty ac id content o f a vegetable o i l feedstock. H o w e v e r , fo r its true potent ia l to be rea l i zed , the l im i ta t ions to t ransester i f icat ion assoc iated w i t h this catalyst must be revea led and ove rcome . Future research w i l l be d i rected toward this goa l . 4.3 Recommendations B a s e d o n the w o r k conduc ted fo r this thesis, a number o f recommenda t ions are p roposed to fo r future research. • T h e es t imat ion o f the A n t o i n e ' s coef f i c ien ts i n H Y S Y S needs to be i m p r o v e d . It is therefore suggested that exper iments des igned to measure the vapour pressure o f t r io le in (or vegetable o i l ) be conduc ted at the temperature range o f interest, be tween 2 5 ° C to 4 0 0 ° C . T h e A n t o i n e coef f ic ients can then be obta ined by regress ing the data, and then input in to the process s imu la t ions . • Expe r imen t s shou ld a lso be per fo rmed to ver i f y that the 3-phase separator used in Processes HI and I V t o remove g l yce ro l can ach ieve the results o f the s imu la t ions . • T h e s imu la ted feedstocks c o u l d be expanded to i nc l ude those w i th F F A contents greater than 5 wt.%, as i n the case o f y e l l o w grease, and h i g h water contents. S u c h factors m a y change the re la t ive e c o n o m i c order o f the processes. • S i n c e the heterogeneous process ind ica ted such p r o m i s i n g resul ts, i t w o u l d a lso be o f interest to conduc t a more deta i led s imu la t i on , where more care is taken to op t im i ze the d is t i l l a t ion c o l u m n s . It w o u l d a lso be des i rab le to i nc l ude k ine t i c i n fo rmat ion (i.e., the effects o f temperature and res idence t ime) i n the reactor m o d e l l i n g to g i ve a more real is t ic representat ion o f the system. • It is a lso r e c o m m e n d e d that the reasons for the fa i lu re o f S n O to cata lyze any react ion shou ld be invest igated, as w e l l as w h y the method o f A b r e u et a l . (2005) fa i l ed . S i n c e the A T R O R o f the heterogeneous process drops d ramat i ca l l y f r o m 5 4 % to 2 4 % w h e n 72 sul fated z i r c o n i a is used, it be e c o n o m i c a l l y advantageous to use the S n O ca ta lyzed process. W i t h respect to the su l fonated char , it is very impor tant to understand and o v e r c o m e the l im i ta t ions to t ransester i f icat ion associated w i t h the catalyst and exper iments shou ld be des igned to e luc idate the p rob lems . • Synthes is o f the catalyst w i t h f u m i n g su l fu r ic ac id has been s h o w n to increase the total ac id i ty ( T o d a et a l . 2005) , w h i c h may have an ef fect on the react ion. It is r e c o m m e n d e d that cata ly t ic tr ials be conduc ted w i t h char treated w i t h f u m i n g su l fu r ic ac id . • T h e char u t i l i zed i n this study exh ib i ted a h i gh l y c o m p l e x , i r regular ne twork o f pores and f ib rous channe ls , w h i c h may pose mass internal mass transfer l im i ta t ions o n the large t r ig lycer ide mo lecu les . C h a r obta ined f r o m the py ro l ys i s o f coa l has s h o w n a less convo lu ted , h i gh l y regular porous structure ( Y u et a l . 2004) w h e n compared to the char u t i l i zed i n this study. Tes t i ng o f a catalyst de r i ved f r o m coa l char c o u l d y i e l d some ins ight in to any resistance that mass transfer m igh t p lay . • A l te rna te ly , it is poss ib le that the l ack o f catalyst t ransester i f icat ion act iv i ty is due to externa l mass transfer l im i ta t ions . If e thanol b inds to the ac t ive sites (perhaps through hydrogen bond ing ) , that c o u l d prevent the non-po la r t r ig lycer ides f r o m access ing the ac t ive sites and be ing protonated by the ac id groups. In the case o f F F A conve rs i on , the po la r c a r b o x y l i c ac id group o f the F F A c o u l d s t i l l access the act ive site, a l l o w i n g the react ion to occur . U s i n g a coso lven t to create one o i l / a l c o h o l phase c o u l d ove rcome any potent ia l external mass transfer l im i ta t ions , and therefore a l l o w the react ion to occur . • T o determine i f reac t ion temperature is the l im i ta t i on , it is r e c o m m e n d e d to test the react ion w i t h the catalyst at e levated temperatures. • A number o f studies to exam ine the per fo rmance o f the su l fonated char i n the ester i f ica t ion react ion w o u l d a lso be usefu l : invest iga t ing the effect o f feedstock water content o n the reac t ion ; de te rm in ing the m a x i m u m amount o f F F A s i n the feedstock before the react ion is i nh ib i t ed ; and test ing the cata lys t 's reusabi l i ty . • Las t l y , de terminat ion o f k ine t i c parameters fo r the ester i f icat ion and t ransester i f icat ion react ions w o u l d be very va luab le . S u c h data w o u l d enhance the s imu la t i on fo r P rocess U l , and be very impor tant fo r the sca le-up o f the react ion to an indust r ia l scale. 73 4.4 References A b r e u , F . R. , A l v e s , M . B . , M a c e d o , C . C . S . , Z a r a , L . F. and Suarez , P . A . Z . (2005) . N e w mul t i -phase cata ly t ic systems based o n t in c o m p o u n d s act ive fo r vegetable o i l t ransester i f icat ion react ion. Journa l o f M o l e c u l a r Ca ta l ys i s A : C h e m i c a l 227(1-2) : 2 6 3 - 267 . Fu j i t a , K . , N a k a m u r a , C , M a t s u d a , K . and M i t s u z a w a , S . (1990) . Prepara t ion o f tin(JI) o x i d e b y a homogeneous prec ip i ta t ion method . B u l l e t i n o f the C h e m i c a l Soc ie ty o f Japan 63(9) : 2718 -2720 . Fu ru ta , S . , M a t s u h a s h i , H . and A r a t a , K . (2004) . B i o d i e s e l fue l p roduc t i on w i t h so l i d superac id cata lys is in f i x e d bed reactor under a tmospher ic pressure. Ca ta l ys i s C o m m u n i c a t i o n s 5(12) : 7 2 1 - 7 2 3 . J i tput t i , J . , K i t i y a n a n , B . , Rangsunv ig i t , P . , B u n y a k i a t , K . , A t tana tho , L . and Jenvan i tpan jaku l , P . (2006) . Transester i f i ca t ion o f c rude p a l m kerne l o i l and crude coconu t o i l by di f ferent so l i d catalysts. C h e m i c a l E n g i n e e r i n g Jou rna l 116(1): 61 -66 . K i s s , A . A . , D i m i a n , A . C . and Ro thenberg , G . (2006). S o l i d ac id catalysts fo r b iod iese l p roduc t i on - towards susta inable energy. A d v a n c e d Synthes is & Ca ta l ys i s 348(1 + 2): 7 5 - 8 1 . L o p e z , D . E . , G o o d w i n , J . G . , B r u c e , D . A . and Lo te ro , E . (2005) . Transester i f i ca t ion o f t r iacet in w i th methano l on so l i d ac id and base catalysts. A p p l i e d Ca ta l ys i s , A : G e n e r a l 295(2) : 97 -105 . Suppes , G . J . , D a s a r i , M . A . , D o s k o c i l , E . J . , M a n k i d y , P . J . and G o f f , M . J . (2004) . Transester i f i ca t ion o f soybean o i l w i th zeo l i te and meta l catalysts. A p p l i e d Ca ta l ys i s a- Gene ra l 257(2) : 2 1 3 - 2 2 3 . T o d a , M . , T a k a g a k i , A . , O k a m u r a , M . , K o n d o , J . N . , H a y a s h i , S . , D o m e n , K . and H a r a , M . (2005) . G r e e n chemis t ry : B i o d i e s e l made w i t h sugar catalyst . Na tu re 438(7065) : 178. Y u , J . L . , Ha r r i s , D . , L u c a s , J . , Rober ts , D . , W u , H . W . and W a l l , T . (2004) . E f fec t o f pressure on char fo rmat ion du r i ng py ro l ys i s o f pu l ve r i zed c o a l . E n e r g y & Fue ls 18(5): 1346- 1353. 7 4

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