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UIC The First Address in Short Path Distillation

Production of High Concentrated Monoglyceride by Willi Fischer Lecture given on occasion of the DGF-Symposium in Magdeburg / Germany in October 1998

Production of High Concentrated Monoglyceride Willi Fischer, UIC GmbH, Alzenau Hörstein

Main consumer of Monoglyceride is the food industry and quantities required are so high, that worldwide several companies have specialized on Monoglyceride as their main production. Monoesters with C16/C18 acid groups are preferred. If a high amount of unsaturated acids is present, e.g. if the iodine value is >80 the final product is a liquid or a jelly at room temperature. If the iodine value is 90 % Monoglyceride with low content of free Acids and Diglycerine and free of Ash is required. For the production of this so called ″High Mono″ the distillation of Low Mono is necessary. This can be done only by Short Path Distillation at approx. 200°C and 0.01 mbar or less.

Pic. 5 shows the sectional view of a distillation unit for a feed rate of 13 000 mt/a Low Mono. The quantity of High Mono recovered from the Low Mono can vary from 40 – 55 c\daten\ausarb\publikat\monovort.doc

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% of the total content depending on what concentration is required. Monoglyceride is the distillate, Di- and Triglyceride remain in the residue. In the Monoglyceride molecule the acid group is coupled with Glycerin either in α- or in β position. The α-Mono content is wanted as high as possible as it has a much higher emulsifying effect as β-Mono. Just after reaction, the β-Mono portion can be 7 % of total. During storage some β-Mono will be converted into α-Mono and the content can be as low as 2 % if conversion is completed. Therefore the specification of α-Mono content is not precise and is depending on the time when the sample is analyzed. In elder literature we find often the specification of α-Mono while in our time total Mono content is specified. Reason is, that the wet analysis which was used in former time can determine only α-Mono while the GC method now used measures the total Mono content.

The residue from the Mono concentrating is recycled to the reaction process and is further converted into Monoglyceride. Approximately 50 % of the recycled residue is converted into Monoglyceride at each recycle. After 6 recycles, about 98 % of the fat is converted finally into Monoglyceride. As for each reaction batch fresh Fat and fresh Glycerin is added the averaged residence time of the material in the process is between 2 and 3 passes. As longer the residence time of the material in the process as more thermal degradation can happen. The acceptable number of recycles is determined by the quality specs. of the final product and is influenced by the quality of the feed material, but more by adequate equipment, precise operating conditions and proper storage. E.g. nitrogen blanketing of material at elevated temperature shall be considered. Manufacturers told as that recycle rates ranging from 3 to more than 25 recycles are usual. Unlimited number of recycle is only theoretically a 100 % use of material. A small amount of Glycerides is purged with the settled Glycerin after reaction. From the material balance of an existing plant over an extended time of operation and with more than 20 recycles, the lost of fat was calculated in the range of 2-3 %.

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High excess quantity of Glycerin in the feed mixture is required for several reasons. It was already mentioned, that excess Glycerin pushes the equilibrium to the right side of the reaction formula. As the catalyst is promoting also the reaction backwards there is a risk, that Monoglyceride is disproportioned during Glycerin stripping Therefore it is necessary to deactivate the catalyst. The deactivation is done by adding Phosphoric Acid. The Phosphate will be extracted from the mixture by the insoluble portion of Glycerin. It is removed from the process together with the settled Glycerin. To avoid recycling of the Phosphate in the process, the settled Glycerin must be distilled before it can be recycled to the reactor.

An alternative which does not require a redistillation of the settled Glycerin each time is the use of Calcium soap as a catalyst. Calciumphosphate is insoluble in Glycerin and can be separated by filtration or with a centrifuge. However this is somewhat difficult and there is a risk, that the ash content is higher in the final product. Also it has to be considered that Diglycerine together with undistilled Glycerin is recycled and will increase the content in the final High Mono.

On production scale, direct Esterification as well as Interesterification can be done continuously or batchwise. As today more than 90 % of Monoglyceride is produced by Interesterification only this route shall be further discussed.

For continuous reaction one equipment configuration can be a mixing vessel followed by a pipe reactor.

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Reflux condenser

Reflux condenser

Steam ejectors

Deactivating chemicals

Catalyst

Cooling water

Reactor 1

M

Steam

Reactor 2

Cooling water

M

Scrubber

Decanter

Cooling water

Liquid ring pump Cooling water

Steam

To hotwell

Fat, Glycerine, Recyclemixture Precipitated Glycerine

Glyceride mixture

Batch Reactor System for Monoglyceride Synthesis Equipment Diagram

Pic.6 Monovort.xls

A simplified equipment diagram of a 3 stage batch reactor system is shown in Pic. 6. Reactor 1 is a steam heated mixing vessel with a reflux condenser. In this vessel the mixture is heated till the required concentration of Monoglyceride is reached. Then the mixture is pumped into the second reactor and cooled as fast as possible, to minimize reverse reaction. After catalyst deactivating the and settling is finished, Glycerin and Glyceride mixture are pumped to separate storage tanks.

The total batch time is almost 9 hrs. However, in a 3 vessel reactor system as shown the vessels are operated in an overlapping mode which allows to finish all 3,5 hrs an other batch and nearly 7 batches can be made in 24 hrs. The batch volume has to be considered carefully. Too small reactors cause higher specific manufacturing and operating cost. Too large reactors designed with a small l/d ratio are expensive again and very powerful, high energy consuming agitators are required. If l/d is high, the much heavier Glycerin phase is not carried quickly enough in the upper section and emulsion equilibrium is not reached.

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Experience taught us, that reactors for a batch volume of 5 - 10 m³ are somewhat an optimum. The design of the reactor system must also consider GMP rules and especially low thermal degradation. Heating with steam rather than with hot oil reduces temperature overshooting. Recommended material of construction is Inox according to German code 1.4571 or American code 316.

Condenser Steam ejectors Cooling water

Stripped Glycerine

Glycerine stripper

Rectfication tower

Steam

Monoglyceride distillator

M

Oil booster pump

Scrubber

Hot oil

M

Liquid ring pump

Cooling

Hot oil

Cooling water

To hotwell

Cold trap Glyceride Mixture after Reaction

High Monoglyceride Quench cooler

Cooling water

Low Monoglyceride

Diglyceride

Continuous Distillation System for High and Low Monoglyceride Equipment Diagram

Pic.7 Monovort.xls

Pic. 7 shows the simplified equipment diagram of a double stage distillation plant for Monoglyceride. In the first stage Glycerin is stripped in a thin film evaporator with rectification tower and an external condenser. Glycerin free Low Mono can be discharged via a quench cooler from this step. For High Mono production, the bottom product is pumped into the following Short Path distillator with internal condenser, where Monoglyceride is concentrated in the distillate.

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Both evaporators are thin film evaporators indirect heated with thermal oil. Vacuum is produced by an oil booster pump followed by a multiple stage steam ejector system and a liquid ring pump. Driving fluid for the oil booster pump and the liquid ring pump is Glycerin. This will avoid contamination with a non food grade auxiliary medium in case of operating or plant failure.

Thin Film Evaporator Series RF

Short Path Distillator Series KD

Pic. 8 shows the principle of thin film evaporators with and without internal condenser. For economical reasons high availability of the distillation is essential. Modern plants are operated with net distillation time of 8000 hrs/a. This is an uninterrupted operation 24 hours a day for 7 days a week all year round. Only one shut down per year for maintenance is necessary. To maintain the high quality specs for the final product, fouling on the evaporator surface must be avoided. Experienced manufacturers of agitated film evaporators deliver equipment with fine polished surface and optimal roller wiper system to fulfill these requirements.

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The economical figures, especially the production of high concentrated Monoglyceride which requires a higher investment shall be analyzed. Pic. 9 shows the annual key data of 1994 of a production plant with a capacity of 10 000 mt/a.

Input Unit price Hardened Soy Oil Glycerine

Quantity

Annual cost US$

930

US$/mt

8.319

mt/a

7.736.670

48,8

2.050

US$/mt

1.815

mt/a

3.720.750

23,5

105.000

0,7

1.400.000

8,8

Rework of settled Glycerine Share of plant cost per year

Portion %

1)

Steam 40 barg

61

US$/mt

8.900

mt/a

542.900

3,4

Steam 12 barg

56

US$/mt

2.400

mt/a

134.400

0,8

305 US$/MWh

750

mt/a

228.750

1,4

880 MWh/a

72.160

0,5

150.000

0,9

260.000

1,6

1.500.000

9,5

15.850.630

100,0

Fuel Electricity

82

US$/mt

Other utilities, instrument air, nitrogen, cooling water, waste disposal, etc. Labor cost

50

US$/h

Overheads Total

2)

Output Unit price

Quantity

Annual gross profit US$

High concentrated Mono 1)

2.140

US$/mt

9.840

mt

21.057.600

Total Plant Investment is 7.000.000 US$. 20% of this sum covers depreciation, intrest rate, insurance, etc. per year.

2)

Cost do not include packing, transport and distribution and sales activity.

Cost/Profit Break Out of Figures of 1994 Plant Capacity: 10.000 mt/a High Mono

c\daten\ausarb\publikat\monovort.doc

Pic.9 Monovort.xls

12

This plant makes good profit even if the additional distribution costs are considered. Interesting is a look on the cost increments. Feed material is almost 72 % of the total cost, while 9% is overheads. Only the remaining 19 % of the cost can be influenced by engineering and production management. Energy, maintenance, labor are small portions and must not be analyzed furthermore. The remaining portion for plant investment with less than 9 % is reasonable.

However as the investment is high, purchase departments tend to decide more on price rather than on technical performance to save money. However this is not smart at all. Due to strong competition amongst bidders for this type of plants the range for high and low price bids are max. 10 %, which is only 1 % cost saving. If less investment would reduce the yield by only 2 % this would increase the feed material cost portion by about 1.5 %. This shows the relation. The same is true for reduced maintenance efforts.

Another aspect is the minimum size where profit is made at all. As some cost factors are constant, others linear and some reverse proportional, there is a minimum production where break even is reached. This is dependent on market situation, infrastructure, environmental regulations, annual load etc.

To support the decision of potential investors, our company has developed an in-house calculating program for quick calculation of economical key data. Pic. 10 shows several cases of investigation of different sizes often requested in an early stage of investment decision. It is astonishing, how much break even can differ depending on the conditions.

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P ro d uc tio n c a p a c ity o f hig h c o nc e ntra te d M o no g ly c e rid e

P ro d uc tio n c o s t re la tiv e to g ro s s s a le s p ric e

m t/a

%

C ase 1 1 .3 0 0 5 .0 0 0 1 5 .0 0 0 1 .8 5 0

107 87 79 (B re a k e v e n)

1 .3 0 0 5 .0 0 0 1 5 .0 0 0 1 .2 3 0

98 81 70 (B re a k e v e n)

1 .3 0 0 5 .0 0 0 1 5 .0 0 0 3 .5 0 0

117 95 86 (B re a k e v e n)

1 .3 0 0 5 .0 0 0 1 5 .0 0 0 2 .0 3 0

108 87 74 (B re a k e v e n)

C ase 2

C ase 3

C ase 4

Case N e t. p ro d uc tio n tim e /y e a r S o urc e fo r ra w m a te ria l F a b ric a tio n p ro g ra m

Infra s truc ture E nv iro nm e ta l re g ula tio ns R e a c tio n p ro c e s s N o . o f R e c y c le s

1

2 7 .5 0 0

O uts id e M o no g ly c e rid e o nly M e d ium le v e l H ig h B a tc hw is e >6

3 8 .0 0 0

Inho us e O the r O le o c he m ic a ls a s w e ll H ig h le v e l H ig h B a tc hw is e >6

4 6 .4 0 0

O uts id e O the r O le o c he m ic a ls a s w e ll M e d ium le v e l Low B a tc hw is e >6

8 .0 0 0 O uts id e M o no g ly c e rid e o nly H ig h le v e l H ig h C o ntinuo us >6

In flu en ce o f P la n t S ize o n P ro d u ctio n C o st C o m p a riso n o f d iffe re n t S c e n a rio s

P ic .1 0 M on ovort.xls

Summarized we can say, that the technology for the production of Monoglyceride has reached a high level so far. Under the pressure of global competition plant manufacturers as well as producers are forced to optimize the technology further. Based on today’s experience remarkable improvement of the distillation technique is not expected. However improvement of the synthesis and optimized operating conditions seem to have room for further improvement in the near future. It will be interesting to see the results.

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