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14/12/2020
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Q:
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Here are a few points regarding the desalter issue in our refinery:
1. There are two desalter trains installed in parallel.
2. When the desalter current fluctuates, the brine water coming out of the desalter has an opaque black colour.
3. The events occur only at one train. In this case, the rest of the trains are fine with the same crude. 4. The events tend to occur at crude switch or introduction of slop oil, but this is uncertain.
What are the possible causes? Please advise what we should investigate and analyze to clarify the causes.
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(4)
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19/09/2020
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Q:
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We are operating an aromatic recovery unit producing benzene and toluene. The extraction section uses Sulfolane as solvent. The extract is stored in a charge day tank and is used for charging the benzene column. To remove olefins from the feed, there is a clay tower prior to the benzene column that operates at at inlet temp of 170 deg C and a pressure of 13 kg/cm2. There is an exchanger for heating the clay tower feed (tube side). We are observing a frequent issue of plugging in this exchanger. This leads us to shut down the fractionation section for almost a day every five months for cleaning/replacing the tube bundle. The olefin content in the light reformate feed varies between 5% and 7%. Is there any way this issue can be resolved? Is the olefinic content in the feed too high? The plugging material seems black in colour. Is there any method that can be used for identifying the fouling type? Is is it due to polymerisation of olefins? Any solution to avoid such frequent plugging in this exchanger?
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(10)
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09/09/2020
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Q:
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We have a sour fuel gas amine absorber where sour fuel gas from the refinery is treated with amine from the amine unit to strip out H2S. Sweet FG from the outlet of the absorber passes through a cooler then a filter coalescer to separate carryover amine from fuel gas. We are continuously getting water from the filter coalescer instead of amine. We have checked the cooler for tube leakage but no leak is observed, also the pressure of the fuel gas side is 0.5 - 1kg/cm2 higher than the cooling water. After checking the strength of a coalescer boot sample, almost 99% water is found. Is this a normal outcome or what are the probable causes of water formation instead of carryover amine?
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(5)
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19/07/2020
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Q:
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How effective is hydrogen in stripping H2S from hydrotreated naphtha?
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(6)
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26/06/2020
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Q:
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In a multi-bed filter why is the larger particle size bed placed over the smaller particle size bed?
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(1)
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08/05/2020
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Q:
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How can I detect my reboiler leak from Column parameters? I have observed no pressure fluctuations during the leak. Why did the column pressure not increase to steam pressure?
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(1)
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15/03/2020
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Q:
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We are operating a Naphtha Hydrotreater with two reactors. The first reactor is for diolefin saturation. We are facing high DP issues in the 2nd reactor. What could be the cause?
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(2)
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24/05/2019
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Q:
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We have a problem at our CDU. It is observed that when we examine salt content at the start of crude charging tank after giving it 24 hour settling time, salts are in yhe range of 2 to 5 PTB. But when same tank is charged and sample is taken at CDU upstream, salt content is higher... Up to 50bPTB... What could be possible reasons?
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(7)
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12/04/2019
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Q:
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In the our HDS unit it treat light distillate cut (Ibp 40 C Ebp 230 C) come to unit from CDU unit and after treated go again to CDU unit to separate the light distillate to products LSR , NAPHTHA and KEROSENE .
In the last 3 weeks till now we had tested Naphtha in the lab. for sulphur content and the result is > 1 ppm the normal result is < 0.5 ppm which is suitable for SR reforming unit as feed ...we do many changes in the operations setting in the stripper column but that change nothing. The stripper column bottom temperature maintained by circulating to furnace and we noted that the level in reflux drum increased abnormally in the last month so we increased the reflux flow rate to the overhead. Are there any recommendations for this situation?
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(5)
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18/01/2019
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Q:
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I have a problem with time consumed in purification through distillation. My distillation column consists of 27 feet 12" column with paul rings as packing components.
I have a reflux at 17 feet in column, that comes from the two condensers of 17 meter square and 10 meter square. I do distillation by two stage water ring vacuum pump with booster attached. I get the vacuum of 735mm/Hg.
What should be the reflux temperature that goes back in column? Or in reverse, what should be the temperature difference that goes back in column and at the bottom?
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09/01/2019
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Q:
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In our VDU column we are planning to route slop recycle directly from Chimney tray (Above flash-zone) to stripping section using gravity flow. Is it feasible to route slop recycle without slop quench pot & pump?
Situated in India
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(2)
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02/01/2019
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Q:
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On our dehydrator vessel we have 2 x 2 phase transformers but the current is unstable. Why would the current be fluctuating so much?
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(1)
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28/12/2018
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Q:
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We are operating a small refinery processing sweet crude (less than 0.4 wt % sulphur). The crude is heated in a heat exchanger network and sent to a preflash column. The overhead from preflash column are condensed as naphtha and sent for stabilization after removing free water in overhead reflux receiver boot followed by coalescer. The naphtha is reboiled in the column and refluxed by a overhead stab in condenser. Vapour from the column are sent as fuel.
Recently when the column was opened up after one year of service the overhead condenser was badly corroded. In fact all the tubes had holes (condenser uses cooling water in the tubes). The strange thing which was noted that elemental sulphur embedded in the corrosion product covering the outside of tubes.
We are wondering where this elemental sulphur was formed? The overhead operating temperature is 100°F.
We are using antifouling agent in our crude but the vendor says that there is no possibility of elemental sulphur from their product.
Additional:
1. Preflash overhead goes through a prefilter followed by a sand bed coalescer. We have observed no emulsion and water haze after these filters and coalescers. However, we are recycling boot water to overhead condenser in the preflash. There is no water wash in the stabilizer as it is a simpler stripper with no overhead condenser and drum.
2. No outside naphtha is being processed; however, demin water solution is prepared with neutralizer which is injected in preflash overhead. We are wondering about this Claus type reaction that take place under these mild conditions without catalyst.
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(2)
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13/12/2018
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Q:
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Reprocessing slop oil is always a headache issue for refiners, in our refinery, we blend recovery oil (one of slop oil from crude tank cleaning) into heavy crude slates (API<28), but suffering problem on desalter operation, higher emulsion layer led to electric field trip and desalter brines contain oil which can result wastewater treatment plant in upset, we would like you could share the operating experience on reprocessing slop oil with minimal impact on these facilities, thank you~
Questions
(1). What kind of pre-treating methodology did you apply on slop oil before reprocessing? Water separation or any filtration steps?
(2). Did you use chemicals for slop oil pretreatment?
(3). How did you reprocess slop oil? In-line injection into CDU uint or blend slop oil into crude tank?
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(6)
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06/12/2018
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Q:
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Crude oil desalter problem:
It is observed that during normal running of Crude oil desalter(2 stage in series), the Amperage increased from 45 to 90Amp. It was also checked that there is no water shot with crude(i.e. < 500PPM H2O). Immediately wash water stopped, still current doesen't comes down. Hence, wash water resumed and observed current at higher side. Evev, the crude oil type processed is also the same as earlier. What could be the reason for high current and suggest solution to bring down current?
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(4)
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31/10/2018
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Q:
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In VDU we are facing problem in HVGO SECTION, that HVGO pump suction strainer periodically choking?
What’s the causes?
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(6)
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19/08/2018
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Q:
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Usually, Steam is used for CDU's stripping. I wonder that can I use fuel gas that comes out of the top of the column for stripping? (& Recirculating fuel gas). I think that by using fuel gas as stripping medium, we can save money & there will be less corrosion at the top of the column.
Can you tell me the advantages & disadvantages about this idea?
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(3)
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17/08/2018
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Q:
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In most cases, steam for stripping at the Crude distillation units. IF I use FUEL GAS for stripping at a Crude distillation unit ,what are the disadvantages? I think it would be nice to reduce the amount of steam used and to gain the economic benefits of it, since it is to recirculate the Fuel gas that comes out from CDU.
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21/07/2018
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Q:
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What back flush procedures are available for cleaning a heat exchanger?
Feed is Meta and ortho xylene on cold side and eulibrium conc of xylene on others.
What are the suggested ways for packinox online cleaning procedures?
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17/07/2018
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Q:
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In our Refinery, we are facing continuous emulsion and water carry over from wash water tower. What all can be the reasons?
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(6)
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28/06/2018
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Q:
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Currently, I am trying to reduce the sulfur concentration from the hydro-treated naphtha. After reading up a few articles I came to the conclusion that the sulfur concentration is due to improper stripping of H2S from the stripper column. I have to improve the performance of the stripper column to reduce the sulfur concentration by adjusting pressure and R/F.
How do I proceed? Is there any other sources of sulfur that I have to pay attention to?
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(6)
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31/05/2018
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Q:
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On the top of a flare tower, there are certain "star" shaped objects surrounding the top. What are the uses of those?
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(1)
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08/04/2018
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Q:
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I have difficulty in drawing product from my side stripper of the atmospheric distillation tower. Whenever I raise the Stripping steam rate, this problem will occur. My initial suspect is due to the hydraulic limitation when the stripping steam is above a certain value. The technical reasoning would be when there is high vapor rate rising up the stripper tower, the vapor load creates high pressure drop across the stripping trays. Liquid flowing from the top will ultimate be restricted from flowing down the stripper tower and creates hydraulic limitation. Do you all agree on this observation?
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(3)
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10/03/2018
|
Q:
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At Temperature above 370'C in a CDU? What cracks?
- Diesel? (Asked b/c FBP of Diesel is 370'c)
- Ends heavier than Diesel?
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03/03/2018
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Q:
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Temperature of the CDU feed is always less than 370'C because temperature further than that will cause cracking. Cracking of what? Cracking of Diesel? Or cracking of heavier ends? I'm asking because FBP of Diesel is 370'C.
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(4)
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14/01/2018
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Q:
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How can we calculate the draw off temperatures of Kerosene and diesel to be with drawn from CDU? Similarly what would be the procedure in case of LVGO and HVGO in case of Vacuum Distillation?
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(1)
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29/08/2017
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Q:
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What is the best way to re-process Fuel Oil Blend Stock (FOBS) or waste that contains 20-30% water and sediments?
Any recommended third party company doing this job?
Any way to sell such material?
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(3)
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24/07/2017
|
Q:
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Is a microbiological assay for storage tanks necessary?
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(3)
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16/05/2017
|
Q:
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During the drying of fusel oil, what are the correct proportions and quantities of fusel oil and sodium chloride solution used for mixing? What amount of the filtered solution is distilled?
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08/05/2017
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Q:
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Fusel oil removed during distillation. It contains various higher alcohols. What is the standard procedure for separating isoamyl alcohol?
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(1)
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12/04/2017
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Q:
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In my CDU unit, there are two type of feedstocks -- sour crude and sour condensate. I noticed both Units have same configurations -- except CDU have desalters and charge heater while Condensate Fractionation Unit does not have them. While crude feed is vaporized up to 60% before charged into CDU column (360degC), condensate feed is heat up only up to 140degC where it is still 100% liquid phase.
My question is,
1) Why in CFU configuration, it does not requires Charge Heater at upstream of Condensate Fractionation column?
2) What is the factor determining the vaporization rate of condensate/ crude feed into the fractionation column?
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(1)
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04/04/2017
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Q:
|
In a vacuum distillation unit working with a multi-stage steam - ejectors system, and it's around 10000 BPD capacity.
The flowing of the suction fluid (hydrocarbon mixture) is 24975 M3/hr and 79 C˚.
Can we switch to vacuum pumps instead of the steam - ejector, from an economical point of view?
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(3)
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26/03/2017
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Q:
|
I am working in Crude distillation unit. Heavy Naptha yield is coming very low. Around 4-5 M3/Hr against 26 M3/Hr. It is observed, whenever HN stripping steam reducing stripper level is increasing. Stripping steam reduced from 400 Kg/Hr to 200 Kg/Hr. Even then we are able to get 4 M3/Hr. Could any one advise to overcome this problem?
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(6)
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15/02/2017
|
Q:
|
In a CDU overhead drum reflux, which are the advantages of a three phase separator versus a flooded one with a naphtha + gas outlet, a naphtha to reflux outlet and a water separation? And how can you estimate them? In both cases the reflux temperature is the same.
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(1)
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19/01/2017
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Q:
|
We have a 50,000 bbl/d capacity crude unit designed for Iranian light crude oil. The main crude column needs to be replaced due to ageing. We would like to take this opportunity to revamp to unit capacity as well to about 70,000 bbl/d. Based on a previous study carried out, the unit capacity can be increased up to 70,000 bbl/d by installing a pre-flash drum before the charge heater. However, now we have to replace the main column. In another study carried out, it has been identified that the some modifications are required to be done to the charge heater such as re-tubing with different metallurgy and changing the passes from 1 to 2 etc. if the unit capacity is increased up to 70,000 bbl/d (without a pre flash drum).
I would like to know whether installation of straight 70,000 bbl/d capacity column or installation of same capacity 50,000 bbl/d along with a new flash drum (to avoid charge heater modifications) is more economical.
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(4)
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04/01/2017
|
Q:
|
In my refinery there is a 15 kBPSD LSRG sweetening unit in which caustic washing procedure followed by MEROX oxidation process. In case of feed change scenario, is there any solution in terms of gas condensate sweetening by means of before mentioned facilities? If yes, what are the changes in terms of capacity, chemical consumption, and mercaptan removal efficiency? If there is any revamp, which sections need to be resized?
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(2)
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04/01/2017
|
Q:
|
The most common (in fact just ) application is stripping steam for stripping at the Crude distillation units. Is there any alternative to stripping steam, such as nitrogen, for stripping of steam at a Crude distillation unit? Is there a literature/research? What are the advantages or disadvantages?
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(2)
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12/12/2016
|
Q:
|
In our Naphtha Stabilization unit, feed after preheating leaves the HE through a 10" dia pipe and then immediately split in to two vertical risers of 4" dia and again joins back to a 10" dia pipe before entering the stabilizer. What is the purpose of this risers with reduced dia? In P& ID it is mentioned as two phase flow.
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(2)
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08/10/2016
|
Q:
|
Question is based on Desalter operation.
1. What can be possible reasons of mud in brine but no traces of oil during normal desalter operation? Is desalter paramters temperature, pressure and mixing delta P plays any role in it?
2. What is the role of Pressure and Temperature in desalter operation?
3. On what basis, Transformer KV setting to be changed in desalter?
4. What is the role of Mixing Valve DP in desalter operation and when does it require to be changed?
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(2)
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28/09/2016
|
Q:
|
Could someone explain me the procedure to calculate the diameter of distallation column reflux accumulator?
|
(3)
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12/08/2016
|
Q:
|
I have a question regarding desalter brine quality, as follows: The desalting of same crude is done by two desalters of different geometry (not two stage desalting, two CDUs). I underline that desalting is done efficiently in both units with respect to salt content in desalted crude. The difference is in the content of sulfides in desalter brine. The desalter that has lower sulfide content in brine is more cylindrical, while the another one "tends to be more spherical". The higher sulphide content represents higher load for waste water treatment plant. My opinion is that this behaviour can occur because of longer residence time of oil and wash water in emulsion volume (or the volume ratio of emulsion volume and total desalter volume). I think that perhaps the emuslifier dossage or or delP on mixing valve are higher. Has anybody faced similar issue in the refinery? Any opinion and experience would be helpful.
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(2)
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20/07/2016
|
Q:
|
We have heatless air dryer which already do a change over dessicant few weeks ago.
As per ANSI/ISA-7.0.01-1996 point 5.1 that :
a. the pressure dew point (PDP) shall not exceeding 4 deg C at line pressure.
b. PDP as measured at the dryer shall be least 10 deg C below the minimum temperature which any parts of the instrument air exposed.
Actual Data :
Dew point : -18 deg C (atm dew point)
Pressure outlet air dryer : 6.6 kg/cm2 g.
our lowest ambient temperature for 17 years is 18 deg C
Design Data :
Dew point : -20 deg C (atm dew point)
Pressure outlet air dryer : 8.5 kg/cm2 g.
I am trying to calculate the PDP using dew point calculator and got PDP design is 6.4 and PDP actual is 6.0 deg C.
Is it allowable to use that instrument air quality?
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06/07/2016
|
Q:
|
We are working in a hydrocracking unit and since we start up, we haven't had a good copper strip corrosion data in the light nafta, we have 4a and the best we have achieved consistently is 2b. We lowered the pressure in the main stripper from 125 to 115 psig, increased overhead temperature from 282F to 292F and increased the stripping steam from 8000 lb/h (design) to 9600 lb/h. Also in the debutanizer we have drop the pressure from 160 to 140 ans still we don't have good results. What can we do in order to reach the nafta copper strip corrosion in 1A?
Additional:
Thank you for your answers, I checked the steam and I saw it is 175 psi and 390F, so we are going to heat up more the steam and we are going to try increasing more the flow, but what could happen if I increase it too much, maybe the control valve 100% open and still not get the copper strip corrosion ok?
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(6)
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29/06/2016
|
Q:
|
We have Shell and tube heat exchanger named E-201-11. This E-201-11 is exchanger just before pass heater of furnace of CDU. The service fluid is desalted crude (shell) and vacuum residue (tube) from bottom column.
Shell operating pressure is 19.7 kg/cm2 and tube operating pressere is 25.8 kg/cm2.
Pressure desain shell 30 kg/cm2 and pressure desain tube 36 kg/cm2.
Hydrotest pressure shell 43.5 and tube 37.5 kg/cm2.
What the main consideration of installing TSV at outlet of desalted crude?
Does it because of thermal expansion?
Now we are installing spare exchanger for E-201-11 but the type is plate and frame HE.
The operation mode will be one HE operated and one HE spare/standby.
My questions are :
Do we need to install relief valve at desalted crude outlet of new HE? Can we use 1 TSV for 2 HEs?
What may cause thermal expansion since reduced crude being pumped by bottom CDU?
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(6)
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|
29/06/2016
|
Q:
|
I am working on performance review of sulphur recovery unit (SRU). Recently we conducted a performance test in SRU. In this regard I wish to know the correct method of measurement of sulphur recovery. Some people believe that sulphur recovery should be merely based on measurement of sulphur in feed minus that in incinerator stack whereas I strongly believe that sulphur balance must be carried out which includes sulphur in feed and stack plus the product sulphur recovered in pit/storage. Material balance is a key to any performance test of the plant. Hence I believe that product sulphur measurement in pit is essential for estimating the % sulphur recovery. Kindly confirm if my understanding is correct. Alternatively please advise other methods of sulphur recovery measurement and the most accurate method to be followed.
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(1)
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29/06/2016
|
Q:
|
Can someone advise the method of sulphur recovery in a Performance test of sulphur recovery unit (SRU) in the sense whether % sulphur recovery can be measured by
Sulphur in feed X 100/product sulphur received in sulphur storage? There are some people who believe that sulphur recovery can be estimated by measuring the sulphur in feed (A) and sulphur species in the incinerator stack (B) and simple subtraction of (A)-(B). However I strongly believe that the sulphur balance of the whole plant should be done which includes quantity of sulphur received in storage pit. Please advise. I would be grateful if somebody advises various methods with pros and cons of each and finally the best method practiced in the industry.
|
(1)
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|
29/06/2016
|
Q:
|
We have Shell and tube heat exchanger named E-201-11. This E-201-11 is exchanger just before pass heater of furnace of CDU. The service fluid is desalted crude (shell) and vacuum residue (tube) from bottom column.
Shell operating pressure is 19.7 kg/cm2 and tube operating pressere is 25.8 kg/cm2.
Pressure desain shell 30 kg/cm2 and pressure desain tube 36 kg/cm2.
Hydrotest pressure shell 43.5 and tube 37.5 kg/cm2.
What the main consideration of installing TSV at outlet of desalted crude?
Now, we are installing spare exchanger for E-201-11 but the type is plate and frame HE.
The operation mode will be one HE operated and one HE spare/stadby.
My question is : Do we need to install relief valve at desalted crude outlet of new HE?
|
|
|
28/06/2016
|
Q:
|
Is it possible to remove metals (Fe,Ni,V) from vacuum slop or vacuum residue streams? If yes, how?
|
(3)
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|
15/06/2016
|
Q:
|
We have crude distillation unit and normally operated at 0.3 kg/cm2 top pressure.
Since the change of feed composition which tend to the lighter crude then the operating pressure reach
0.9 kg/cm2.
do you have any experience to limit the operating pressure of CDU in order to make sure valve tray inside the column not released from the tray itself?
More info:
Maximum allowable Working pressure of CDU Column is 3.5 kg/cm2
Desain operating pressure is 0.6 kg/cm2 (top pressure).
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(4)
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|
14/05/2016
|
Q:
|
Are there other methods of removing salt from crude oil besides using Desalter?
|
(3)
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|
11/05/2016
|
Q:
|
LPG amine absorber is design for 15t/h sour LPG from crude unit and delayed coker unit and we are running it at 20t/h. Amine flow rate is 24t/h for absorption. Column operating conditions is 13.8kg/cm2 and 36degC. Delta T between lean amine and LPG is maintain between 5-8degC.
We are facing problem of continuous Hydrocarbon carry over in Rich Amine from absorber. What can be the possible reason?
|
(5)
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|
07/05/2016
|
Q:
|
What is the definition of Overflash in a crude distillation column? What are its advantages and disadvantages? Does it ensure liquid flow between gas oil draw off tray and flash zone?
|
(7)
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|
10/03/2016
|
Q:
|
We are producing bitumen from atmospheric distillation column thanks to low API crude. In crude distillation unit, all atmospheric residue is directly sent to the bitumen tankage. We can produce the 50/70, 70/100 and 160/220 bitumen spec. easily from atmospheric distillation columns.
However, we have some obstacles to get on spec about increase in softening point in bitumen. We have maximized the furnace outlet temp., stripping steam rate and also diesel draw amount but still softening point value is not in limit value.
When we maximize the all operation parameters, increase in softening point is still above the spec values. (Limit value is 9 C for 50/70 bitumen, we could get 11 C which is the best value that we have got.)
At this point, we want to ask a question that do you have any practice about increase in softening point in bitumen? Do you have any advice to get better increase in softening point in bitumen?
|
(1)
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|
03/02/2016
|
Q:
|
We have LPG caustic wash and water wash systems.Similarly,we have Naptha caustic wash and water wash systems.Frequent caustic carry over in product LPG and Failures in copper corrossion due to exhausted caustic solutions is a operational problem.Is there any continuous monitoring instrumentation available to check circulating caustic strength in caustic wash system? Similarly any instrumentation exists for monitoring recirculating wash water for caustic carry over symptom needing wash water replacement?
If so advise/share
|
(4)
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|
11/12/2015
|
Q:
|
The bottom part of our atmospheric distillation tower operates at 350oC, and we are giving the superheated steam to the column at around 420-430oC. Let's say, we want to decrease the temperature of superheated steam to 400oC but still it is superheated steam of course. Might it have any effect on the distillation of the products or the quality of the products to decrease the temperature in this manner?
|
(7)
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|
04/12/2015
|
Q:
|
For the first startup of crude distillation unit, since we did not have any startup gas oil for flushing procedure (cold and hot circulation), we used the crude as flushing stream. According to operating manual, after establishing of gasoil we should store it into relevant storage tank and then stop the process. After that we should flush the system with existed gas oil, again. I want to know, when we are producing gas oil why should we stop the process and go back to flushing mode?
|
(4)
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|
18/10/2015
|
Q:
|
We are looking for a non hydrotreating based technology to decrease condensate sulfur content to lower than 200 ppm. There is a condensate stream in our refinery in which its sulfur content decreases from 3300 ppmw to 1000 ppmw by caustic wash and we need a further decrease of sulfur content to minus 200 ppm, but not with hydro desulfurization. Please advise.
|
(5)
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|
01/09/2015
|
Q:
|
In CCR platformer high purity hydrogen is produced using PSA, PSA outlet tailgas contains 40-50% H2, flow is about 4-5 TPH and pressure 5-6 Kg/cm2g. Can we use this gas to recover further Hydrogen by PSA or membane separation processes? We have a spare PSA available. Can we feed low purity gas to PSA?
|
(2)
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|
14/08/2015
|
Q:
|
Many of the operating plants inside a refinery complex use fresh caustic to sweeten the gases/Naphtha products. Almost all of this spent caustic is then transferred to Effluent Treatment Plant of this spent caustic. Are there processes available which may be utilised to regenerate this spent caustic?
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(4)
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|
15/04/2015
|
Q:
|
Has somebody experience with petroleum hydrocarbon resin simulation and separation?
|
|
|
02/02/2015
|
Q:
|
During Crude distillation unit start up activities, water travels from crude storage tank to crude tower when furnace outlet temperature was 172C. It caused crude tower trays to dislodge. What if level of crude tower remains high then flash zone, does level of crude tower have significant effect on tray dislodge? Our system is furnace operated crude tower.
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(2)
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|
13/11/2014
|
Q:
|
Please could someone explain the difference between the following:
1. Internal reflux
2. Circulating reflux
3. Pumparound
|
(2)
|
|
20/10/2014
|
Q:
|
We are operating Born Canada Vertical furnace in our refinery for crude oil distillation. It is being observed that some tubes comes closer to each other. And some are swinging like pendulum although we are operating at 80% of design flow rate. Can you share reason of this behaviour.
|
(1)
|
|
20/10/2014
|
Q:
|
What quantity of steam is required in distillation column and side strippers per barrel of crude/products.
|
(3)
|
|
15/10/2014
|
Q:
|
Can you please advise some literature sources or design guidelines for Naphtha Stabiliser design.
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(1)
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14/10/2014
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Q:
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Fuel gas knockout drum is sized for minimum surge time of 2 minutes between 19.5 barg to 16.5 barg for compressor change over.
I would like to calculate the surge volume. How?
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(1)
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21/08/2014
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Q:
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If top pressure is maintained on lower side, what is the effect on top temperature, side draw - off temperatures, pressure profile in the column and on the quantity and quality of overhead product?
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(2)
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18/08/2014
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Q:
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In an Reformer Stabilizer Debutanizer Column, we do regular water washing of the column to get rid of the ammonium salts. We do this procedure by reducing the throughput and pressure of the column and produce off-spec reformate during the process.
We do like to ask if any refiners have a practice of introducing steam into the column while the unit is online to clean the ammonium slats deposits in the column and condenser? If yes, what are the concerns and precautions to be observed?
Additional:
I would like to confirm that what you had mentioned. HIGH PH contributing to the severe corrosion. We have a similiar system upstream(the first column for the FRN Feed) and found severe corrosion in the overhead system of the distillation column and we found that the pH was very low and ammonium salts, in the range of 4.5. Hence,we are injecting a highly basic chemical to increase the pH and are currently maintaining 9 pH. But to our confusion , we are still finding a very high amount of corrosion. If what you mentioned is true, what we did in the system is not going to help us but rather worsen the condition?
Thanks Stephan, Could you please elucidate on the corrosion due to high pH? We have a Debutanizer Column , the first column in the Aromatics Complex which is severely corroded in the overhead due to ammonium salts. The feed is from the refinery , Full Range Naphtha. We had initially of an pH of less than 4. Then we injected an chemical to boost the pH and are currently mainly in the range of 9 pH. But the corrosion is still not under control. Could the high pH be one of the concerns to look at?
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(2)
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13/08/2014
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Q:
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Recently we have suffered some problems of Cupper Corrosion test failure in LPG. The LPG came from a caustic treatment for mercaptan sulphur removal. After caustic treatment, the LPG pass through a decanter (with NaOH/MEA solution) and sand filter, which are supposed to remove any caustic carryover from LPG. We do not see any caustic collected in the sand filter, however we have detected Na and nitrogen in LPG, so we suspect that it is not working properly. The sand filter seems not only not working, but also accumulating some contaminants: we have seen sometimes that LPG pass the cupper corrosion test in the inlet, but not in the outlet of the sand filter.
We are evaluating the possibility of substituting the sand by any other more effective adsorbent for caustic / nitrogen (amines). The possibilities are: activated carbon, Anthracite or alumina.
Has anyone experience with adsorbents for contaminant (caustic, amine, etc..) removal in LPG? Any idea / recommendation regarding the operation of the sand filter?
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(2)
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17/07/2014
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Q:
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Need all your expert views on crude oil Basic N2 impact on fouling tendency. This is limiting on crude flex/ optimization as the refinery has CAM limit for basic N2 (150ppm). Need to understand the fouling tendency of Low N2 crude whether this is credible or perceived. Also understand the fouling tendency/reversibility. If credible, please provide if there are ways to mitigate (eg: every low N2 crude processing is followed by crude that can act as cleaning and recover any loss in duty?)
Low basic N2 could be good for LRCCU feed and also hope for HCU where as this limit could restrict such crudes from buying/processing… We always used to be on the basis of waxy vs. asphaltic… every waxy run followed by a aromatic/naphthenic crude run to provide cleaning effect. Antifouling was other alternate only in LR circuit and /or SR circuit.
Blending of crude based on compatibility to mitigate was another option…
There should get some clear guidelines for mitigation if the impact N2 is credible and proven… can you provide any such details and what is minimum technical solution for such mitigations as this will be a clear big lever for crude flexibility.
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(1)
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18/06/2014
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Q:
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Do anybody have experience in treating Brine from desalters through Tricanters to separate Oil and sludge from Brine?
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(2)
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29/04/2014
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Q:
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I'm working on Kerosene hydrotreating unit simulation to remove sulphur from kerosene by using hydrogen. What are the possible components when it mixed?
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(2)
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29/04/2014
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Q:
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We are processing a heavy crude of API 18. Salt content of the crude is 80-100ptb. BS&W is 0.8 to 1.2%.
We are using stripped sour water as wash water, made up by BFW. pH of wash water is between 7-7.5.
We are maintaining a desalter temperature of 150 deg C.
We are having two desalters in series, which is supposed to bring down the salt by 99%. but we reached up to 90% earlier.
Last two months we are having a high emulsion band. The BS&W of desalted crude is 2-3 and the oil in brine is 1-2% even at minimum delta P (0.3kg/cm2). Water injection rate is 3% to individual desalters (circulation not done as oil carryover with brine observed.
The voltage across the grid is as low as 6-10 KV (Tapping at 22KV). We have changed the secondary tapping to 18V and the deslater is showing an improvement in Voltage (10-12 KV)
1.what can be the reasons for this upset?
2. Can you explain the effect of change in secondary voltage?
Additional info:
We are not adding any scavenger to the crude.
Stripped water pH is 6-6.5 and brine water pH is in 7.5 range.
What parameters of crude oil should I check?
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(4)
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20/03/2014
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Q:
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We face the problem of Diesel salt dryer drain line blockage. Anyone faced this issue? What solution did you apply?
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21/02/2014
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Q:
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In the crude distillation unit, we face problem with Gas Oil colour. Any one have any idea to solve this problem or any one have seen like this in any refinery?!
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(8)
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05/02/2014
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Q:
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Is there any SO2 production due to decomposition of Sulfolane use as a solvent for aromatics extraction?
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(2)
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05/02/2014
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Q:
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In our vacuum distillation column three valve trays are replaced from glistch packing in order to obtain more deeper cut of SAE-40 and wash recycle is provided to wash the packing but whenever unit is down due to any failure the wash recycle line gets plugged. We are using SAE-20/SAE-10 as a wash recycle oil. Can we use HVGO for that purpose or any other solution for that problem?
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(3)
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04/02/2014
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Q:
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We are trying to figure out how to improve the feed control to our new Hydrocracking and Hydrotreater Units, since one of the feeds comes from the Coker Unit, we want to know how variable are the quality and flow of the HCGO, Naphtha and LCGO, because we are aware it would be changing while coker cycles are taking place. We don't have tanks to store LCGO and Naphtha as feed to the units, so these streams go to the hydrocracker and hydrotreater directly from the coker stripper, and if there is a sudden change in composition or flow, it could lead on a runaway.
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(3)
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23/01/2014
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Q:
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Does overhead water wash play any role to achieve dew point early?
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(2)
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20/01/2014
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Q:
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I have some queries about UOP OLEX process:
1. What is total feed sulfur limit to Olex unit?
2. What are temp. & press. criteria of this unit?
3. Do we get only linear olefins from this unit or all kinds of olefins?
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10/01/2014
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Q:
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We have a single stage desalter that will run more heavy crude about API=23, Wash Water about 7%, T=280F. There are 2 Transformer about 150 kVa each. Is there a good way to assess what this existing transformer can handle in term of max conductivity of crude ? Are there easy calculations one can do if you had crude conductivity, etc?
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(2)
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10/01/2014
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Q:
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For an existing single stage deesalter operation that plans to run a heavy crude API of 26, wash water 7%. Vessel has two 150 kva transformer. Are there easy ways to see what is the max crude conductivity that this existing desalter grids can handle? Are there any general guidelines to make such assessment.
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16/12/2013
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Q:
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Is there any effect of coil steam to improve vacuum bottom viscosity introduced in a vacuum furnace?
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(1)
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07/12/2013
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Q:
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What the causes of naphtha diluted in water boot of overhead? visually water boot carbonized and foaming.
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(1)
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02/10/2013
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Q:
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Does anyone know where I can get access to a bauxite percolation pilot plant or contract manufacturer to decolorize wax?
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19/08/2013
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Q:
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I have design a liquid distributor for non-foaming system and design liquid load of 5.1 m3/m2.hr and vapor load of 2640 m3/m2.hr. The desired turndown is 50% and turnup is 110%. The column ID is 300mm. The distributor details areas follows:
1. The distributor is 2 - level deck type distributor i.e. instead of punching orifices/holes in the distributor deck, there are pipes fixed to the distributor deck with each pipe having two holes (bottom and top) at a distance of 50mm. This ensures wider operating ranges. distributor deck is circular one piece sheet of 2mm thick
2. The number of drip points (no. of pipes) is 3 and at turndown the liquid head from the bottom orifice is 35mm.
3. The dia of lower orifice is 3 mm and upper orifice is 4mm.
Could someone please advise if the type of distributor selected is proper and the design is adequate to ensure uniform liquid distribution.
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(1)
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13/08/2013
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Q:
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I am studying the conversion of two stages crude dehydration/desalting train to operate as two separate units due to increased wet crude oil production from oilfields. Much apprecaite if someone can share thier experiences, what factors / paramters should be considered for the conversion and is it really feasible to proceed further.
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(2)
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08/07/2013
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Q:
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What is typical vacuum column off gas composition? We operate our Vacuum column at 410 deg C and 55mm Hg top pressure, recently we are getting high concentration of CO (about 40-50 ppm) in seal pot area where off gas condensate is washed.
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05/07/2013
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Q:
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We have a liquid product named HCGO; ideally it's 280-430 cut material. We are analyzing its distillation by D86 method. same liquid sample when tested with D1160 recovery results were different. Since there is huge difference between 350+ recovery points we are confused as to which method to follow.
1. How to compare D86 & D1160 values - which are more accurate?
2. What is the range of D86 & D1160 test methods wrt. recovery points?
Below is table for reference. Both the results are reported up to atmospheric values and in DegC. (OOR = Out of Range)
S. No Distillation D-86 D-1160
1 IBP 287 280
2 5% 339 337
3 10% 347 354
4 30% 363 385
5 50% 374 403
6 70% 384 420
7 85% 396 437
8 90% OOR 446
9 95% OOR 461
10 FBP OOR 497
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(3)
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02/07/2013
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Q:
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Issue : Since commissioning our coker naphtha yield remains always on higher side by 1 to 1.5 wt%. The quality of the Naphtha end point also remains on higher side 145-150 Deg C than the design value of 125-130 Deg C. We are operating our fractionator with top temperature 99 Deg C & pressure of 0.56 Kg/cm2 G. Top temperature, reflux flow rate & pressure are same as design conditions. We tried simulating the scenario but could not get any clues from that.
Queries:
1. What may be the probable causes of deviation in Naptha end point from design?
2. To what extent can we reduce our top temperature, to drop heavy end of Naptha to LCGO cut below?
3. What are concerns foreseen for low fractionator top temperature operation?
4. To what extent Naptha quality degrades if section trays are damaged or reflux distributor is not working properly?
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(3)
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03/04/2013
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Q:
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I'm working on a study to design a new control schematics for Crude Distillation Column Pressure Control. Any ideas for CDU pressure control strategies?
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(1)
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02/03/2013
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Q:
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What is the process of speciality chemicals derived from Kerosene, mainly :
1. aromatic solvents (Mineral Turpentine Oil MTO type) for paints & varnishes, pesticides
2. dearomatised / hydrogenated aliphatic fluids (white spirits) for alcohols, ethers, esters : for perfumes, cleaning agents, etc.
3. paraffinic compounds for foams and dry-cleaning synthetic / woolen clothes
4. heavy distillates as solvents for commercial dyes & inks
5. high boiling solvents as lubricants in metal cutting / rolling industry.
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(2)
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22/12/2012
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Q:
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Is there a new and advance technology or method to recover the dissolved ammonia in waste water in a pure form?
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(2)
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27/11/2012
|
Q:
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Between flooded condenser pressure control and hot vapor bypass control in a distillation column which one is more preferable and under what circumstances?
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(3)
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10/11/2012
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Q:
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We are operating a small refinery in which the crude column has three side draws with h naphtha, kero, diesel there is a kero pa and diesel pa. There was coke in the suction of resid (crude bottoms) pumps to find the cause we opened the coloumn on inspection little coke was found however the trays below kero pumpa were displaced and crumpled no possibility of adding water was observed as stripping steam was added after through purging and was invisible(super heated) steam addition However there was hammering at kero pa on start up which stopped after increasing kero pa temperature to avoid adding subcooled liquid in a sparger on low flow (at start up) the trays above the kero draw were not affected at all
Can any one throw light what are the possibility to look for?
Running the unit again without identifying the cause of accident doesn't make sense. Besides water surge, which doesnt appear to be the cause, is there any other cause?
We were running the column at 30 psig pressure (design operating) is 15 psig can that be a cause? But diff pressure was normal though slightly higher as we were operating with high naphtha yield than normal trays were operating at 50 to 60 npct flood as calculated thru simulation.
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(3)
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08/10/2012
|
Q:
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Is that correct that a desalter can't remove organic salt? If not, why not?
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(3)
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23/09/2012
|
Q:
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What is preferable for me to enter a process water or condensate to crude oil desalter?
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(1)
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17/09/2012
|
Q:
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If caustic dosing suspended due to some unavoidable reasons is it possible to reduce overhead corrosion (caused by hydrochloric acid) by increasing amount of neutralizer like ammonia or amine at overhead of the Atmospheric distillation unit?
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(3)
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07/08/2012
|
Q:
|
Our Desalter transformer has got three transformers supplying to individual grids. This step up transformers are facilitated with three outlet tappings with higher ratings. We always run the desalters with same outlet tappings for all three transformers. Is it advisable to run it with different outlet tappings in all three transformers in a desalter?
I would like to know electrical feasibility as well as process advantages?
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(2)
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12/07/2012
|
Q:
|
In our refinery in India after revamping of Hydrocraker unit for increase of Hydrocraker unit load to by 30% than earlier, Fractionator column perfomance seems not stabilize at all. There was minor modification done in fractionator column for increase of 30% load by replacing some earlier tray with high capacity valve trays & providing some packing ring inside column. But we are now facing problem of high column pressure of around 1.45 Kg/cm2 against design of 1.1 Kg/cm2. For that column remains upset most of the time as naphtha is not removed properly. Our product flash point is also found lower than design. To compensate for flash point & removal of naphtha from product we always kept column top temeprature at higher side of 99 0c than design 93 0c. For that our high naphtha production always remains a concern as light end section found upset. It also observed that our column feed inlet temeprature always kept slightly lower @360 0C than design of 374 0c. Presently we bring down the column pressure to nearly design pressure of 1.1 kg/cm2 by running two offgas compressor. Now I just want to know can we now keep column top temeprature 93 0c as per design for less naptha production and also meet the flash point requirment of products by increasing the column feed inlet to design 374 0c.
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(3)
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30/06/2012
|
Q:
|
I am trying to build a model to optimize the operation of Crude Desalter and study its effect on Crude Column Overhead Corrosion.
The major salts present in Crude are NaCl, MgCl2 & CaCl2; but in our laboratory we measure only Total Salt Content of Crude (before and after Desalter); we do not measure individual salt.
My queries are:
1) How the individual salt affect Desalter performance and Crude Overhead corrosion
2) Is it required to measure the individual salt's content in Crude?
3) Can I assume some typical break-up of individual salt (Note that the type of crude we process changes very often).
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(1)
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19/04/2012
|
Q:
|
MEG regeneration system. In our plant we have 2 rich MEG tanks that receive MEG/Condensate/Water solution from condensate flash vessel. Last time during pigging activities we receive many sludge from offshore, and now all this sludge is settled down inside Rich MEG tanks. MEG Regeneration package performance rapidly reduced, pumps could not deliver Rich MEG to regen, strainers getting clogged very fast, HC compartment of MEG flash vessel in MEG Regen package filling rapidly. Any ideas how to improve situation with Rich MEG tanks? Maybe clean Rich MEG using hydro-cyclones, or any other equipment? Any links to useful equipment to be installed, or to similar problem anywhere? TQVM in advance.
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|
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15/04/2012
|
Q:
|
We need to build very small vacuum distillation unit . We cannot find out how many of oil will crack and we cannot evaluate how many m3 of gases will be generated . So our questions:
What should be a capacity of vacuum pump in m3 per 1t/h ?
How many gases are usually released ?
or give examples from your plants.
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(2)
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28/03/2012
|
Q:
|
In desalting, does the wash water added to dilute the crude get emulsified with the crude after mixing across a mixing valve or does this remain as free water?
Does this wash water have to be considered while we determine the grid KVA requirement i.e. for separation of emulsified water from the crude?
Do we always have to add de-emulsifier chemicals in a desalter or can we operate without them?
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(1)
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28/03/2012
|
Q:
|
I have following questions on desalter:
What is typical salt content at crude out from 2 stage desalter?
Does mixing of crudes result in more salt slip from desalter than design?
Does wash water salt content affect the salt removal efficiency (more salt comes out from desalter)?
Does inadequate/inefficient demulsifier result in more salt slip from desalter than design?
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(4)
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06/01/2012
|
Q:
|
We have a p.d compressor which takes suction from crude tower reflux receiver the gas composition is c3. c2 c1 little contents of c 4 and c5 but the excessive quantity is h2s gas. K.O vessel collect the liquid (gasoline ) which is sour. How we can save and use this gasoline after removing h2s from it? Any method for separating h2s from gasoline?
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(4)
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|
21/12/2011
|
Q:
|
We have a potential gas plant to process gas at 900psig vol is 250 mmscfd with 8% CO2 with no H2S. Can someone advise me what process to use Membrane or Solvent. Will physical solvent be better or DEA, MDEA for producing pipeline quality gas ie 2% CO2.
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|
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18/11/2011
|
Q:
|
Regarding the LPG Sulfrex Unit in RFCC, I have some questions.
We experience the increase of C4 sulfur content last Saturday (11/12) by the forming of the amine absorber(T-20701). ** a brief unit description is bottom of this writing: sulfur content of C4 goes up from 1~3 ppm to 16~18 ppm
Thus we replace the caustic of prewash drum(D-20702) & Extractor(T-20702). But the sulfur content of C4 is not decreased.
Investigating the cause of amine absorber foaming, we find the significant change of amine absorber condition.
First is difference of amine inlet/outlet flow. Inlet lean amine flow is +6~8 m3/hr higher than outlet amine flow in amine absorber.
There is amine carry over to overhead LPG side in amine absorber.
Second is LPG carry under to rich amine side in amine absorber. Rich amine goes with LPG to amine flash drum before amine regenerator.
So the pressure of amine flash drum sometimes rise to almost drum design pressure.
Finally we replace the activated carbon filer in rich amine side, but there is nothing wrong in amine quality.
After that, Inlet and outlet amine flow is same and the delta P of amine absorber increase to normal condition
We wonder why LPG absorber goes back to the normal condition after replacement of rich amine filter.
Q1. Could you explain the reason for this phenomenon?
Q2. If amine quality is main cause, could you recommend the new guide of amine or other countermeasure?
**Brief LPG Sulfrex unit description :
LPG feed from R2R GAS Recovery unit is sent to the Amine absorber(T-20701). Hydrogen sulfide is removed by counter current of amine solution and the LPG leaves the top of the column and flows into the amine settler D-20701 and rich amine is leaves the bottom of the absorber to amine regenerator.
LPG flows into the caustic prewash drum D-20702 for removal the last traces of H2S not removed in the amine absorber.
D-20703 is Caustic Settler. The settler drum allows to separate and return the entrained caustic to the oxidizer T-20703
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(1)
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|
04/06/2011
|
Q:
|
CDU overhead system
To provide a CDU with top reflux there are several configurations. The two main ones are as follows
Configuration 1: Double condensation
Configuration 2: Top pumparound
In configuration 1, total overhead vapors are condensed in two condensers in series. In the first condenser part of vapor equal to the required top reflux is condensed, water is separated and hydrocarbon liquid is returned to column as top reflux. The remained vapors (this includes the top product plus non condensable vapors) is routed to the second condenser.
In configuration 2, the top reflux is provided by a top pumparound and the overhead vapors ( (this includes ONLY the top product plus non condensable vapors) are condensed in a condenser and no liquid is returned to the CDU.
For a refinery with capacity above 100,000 bpd what configuration is recommended? Considering in both case energy is used in top/feed exchangers network. Can Heat Integration cause to use one of these configurations? We know in configuration 2 some more stages are needed as we have added one more pumparound!
Additional:
Thank you for the answers! Don't you think, the double condensation configuration results in lower flowrate in top section? When I can remove water in the first condenser, why not to return the reflux in lower temperature! furthermore I think thermodynamically, configuration 1 is better than configuration 2. As Ralph stated, the latter also needs at least two/three more trays for top pumparound! What I am not sure is energy saving! it is believed that configuration 2 results in a better heat integration.
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(2)
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|
30/05/2011
|
Q:
|
Some steam Jet ejectors are designed with a nozzle extension. What is the role of this extension in the ejector performance? During the last shutdown of our VDU, we noticed that the first (and largest) ejector steam nozzle was mounted without such an extension.
How could this impact on the ejector performance?
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(1)
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|
10/05/2011
|
Q:
|
We are working on a project to eliminate ammonia from sour gas stream to SRU in one of our refineries in order to improve SRU operation and beside to produce ammonium sulphate to be used as fertilizer. Instead double stage SWS, we are thinking in absortion with sulfuric acid in a tower and neutralization of acid excess etc. One way is to develop the process by own but we´d like to know first if someone know who is licensing this process or who is using this technology.
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(4)
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02/05/2011
|
Q:
|
We have a thermosyphon reboiler for Light kerosene stripper with LGO pump around as heating media. We have seen a significant reduction in the reboiler performance observed by reduced delta T across LGO Pump around. This has resulted problems in light kerosene flash point. We introduced stripping steam and somehow overcome the problems. We observed that there was gradual increase in delta T over a period of time across light kerosene side in the reboiler. Can anyone help us to explain this unusual phenomenon in LK reboiler? Can the dryness across light kerosene in the reboiler results higher delta T? Gamma scanning of the column was done and found normal. How can we predict the Thermosyphon Circulation rates? Can the high temperature of LGO pump around inlet to reboiler cause fouling on kerosene side? If it is fouling on the tubes, then there should be reduction in delta T across LK side. In fact we are seeing an increase in delta T across LK side of the reboiler.
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(4)
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07/04/2011
|
Q:
|
We are facing problems with the main fractionator reflux drum bootwater high chlorides(120 ppm)
There is prefractionator ahead of main fractionator but we are getting zero chlorides in overhead boot water.
does the inorganic chlorides dissociate more at temperatures greater than 250 degC which is prefractionator temperature?
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(3)
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04/04/2011
|
Q:
|
For the last couple of weeks, ATF product ex Merox is failing on Silver Strip Corrosion test. What could be the reason?
|
(3)
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23/03/2011
|
Q:
|
We would like to go for absorber to remove water from methyl acetate.
Feed composition: Methyl Acetate: 99% and water 0.75 % and rest are methanol and acetic acid.
I would like to know which type of absorbent I have to choose to absorb water from methyl acetate.
It will be great help, if someone can throw light on this.
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(1)
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|
06/03/2011
|
Q:
|
I am working in Diesel Hydrodesulfarisation Unit. In our unit after H2S removed in Amine Absober, Diesel will go to Stripper. Where steam (direct steam) used as stripping medium. The purpose of steam stripping is only to removal of lower ends or it will remove H2S also. Presently we are maintaining Stripper I/L and bottom temperatures 230 and 225 deg respectively. If I decrease the temperatures by 5 deg, I will gain that heat in preheat, but it is believed that if we decrease stripper IlL temperature H2S in product Disesel will be more and copper Strip corrosion problem will appear. Is it true , By decreasing stripper I/L temperature and increasing steam Can I balance it.
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(4)
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|
19/02/2011
|
Q:
|
What are the fundamentals to separate hydrogen sulphide and ammonia from two stage sour water stripping unit?
|
(3)
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|
22/01/2011
|
Q:
|
To what extent can we blend fuel oil into gas oil without affecting the viscosity characteristics and maintaining the flash point specifications for gas oil or to keep them within the allowed limits?
Additional info:
First of all we don't have neither FCC, Hydrocracker nor VDU...we only run a conventional CDU
the objective here is to maximize the yield of gas oil...(we call it solar in our national markets) by extra stripping out from fuel oil or residue...the question is; Is there any equations or experimental methods to calculate or estimate the resulting viscosity and flash point of either the gas oil or fuel oil?
Thanks a lot.
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(4)
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16/01/2011
|
Q:
|
Pyrolysis gasoline from Ethylene unit is sent to a recovery unit to recover C7 minus components. These are recovered in two columns under vacuum. Maximum temperature is at the bottom of the second column which is ~ 145 deg C. Unrecovered stuff is sent to Utilities as liquid fuel.
Anti-oxidant injection is done in the Ethylene unit as Pygas contains precursors such as dienes which can lead to polymerisation.
Recovery unit was operating steady, without any problems, for 8 months. Now for some reason the frequency of choking of the strainer of bottoms pump of the last column has increased dramatically. Also, we are experiencing frequent choking of burner guns. Material found is coffee coloured granules which become powder when subjected to pressure.
Trying to understand root cause. Not much has changed in terms of operating conditions. Very few component analyses are done in the whole system and not much information is available.
Hope to get some inputs based on experience in similar units.
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(2)
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03/01/2011
|
Q:
|
What are the key factors for improving efficiency of API gravity oil separator?
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|
|
06/12/2010
|
Q:
|
We have a methanol-Water stripping column, which uses direct injection of LP steam for stripping.
I want to know if it is better to use reboiler instead of steam injection.
Is there is any advantage in using direct injection of steam in methanol-Water stripping column?
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(3)
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01/12/2010
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Q:
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How do you calculate grid area for vertical grids (plate/tube type) in case of a horizontal Desalter?
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18/11/2010
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Q:
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There is continuous increase and decrease in our column delta pressure in water methanol column. At the same time we noted that our temperature profile of the bottom and middle bed is also fluctuating.
I feel that our column is having vapor cross channeling.
There is some variation in feed flow and steam flow, but column is somewhat running at 100 % load.
If anybody experienced such problem in your plant, please throw some light to understand what causes this fluctuation in delta pressure and temperature profile in the bed and what action to be taken.
Additional information:
Steam direct injection for stripping
There are three bed made of PP intolox saddels
Steam flow is controlled by mid bed temp
Reflux is controlled by feed flow
More information:
This is a packed distillation column to strip methanol from water. We are using steam stripping in our case because there are some traces of Acetic acid in the bottom. To prevent corrosion we have to strip at low temperature, so we are using steam stripping. There is huge variation in temperature profile of the middle bed, at 100 % load First indication of channeling is the change in delta pressure and disturbance in temperature profile. Disturbance in temperature profile is caused by improper distribution of vapor flow in the bed. So thinking this is because of vapor cross channeling.
If it is channeling or flooding how can we deal with it?
More information:
Thanks a lot for all your suggestions, we have opened our tower found that steam deflector plate was installed wrongly, so steam was injecting directly into the packing, which caused packing to expand and that caused channeling in our tower. After rectifying this, now we don't face this problem.
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(5)
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09/11/2010
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Q:
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Can you please advice what type of corrosion inhibitor, biocide, antifoulant and polyelectrolyte polymer can be used in Desalter effluent?
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(4)
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05/11/2010
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Q:
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Lately we had a problem on one of our old (33 years) but reliable Desalter Transformers. The Transformer is NWL 150KVA 4.16KV/23KV (3each) on top of our crude oil desalter trap. Two of these transformers experienced failure and we found the high voltage cable was burned and cut. The oil in the entrance housing was discolored (black) and the oil in the third transformer was clear. Later on we found the secondary on the third transformer was disconnected. the floating switch and the grid were OK.
What are the possible causes for such failure?
We replaced the entrance bushings with the high their high voltage power cables for the damaged bushings.
We performed Megger test and Polarization Index (PI) test on the transformers. The PI was 1.4, 1.6 and 1.5 for these transformers. Are these PI readings acceptable to put back the system in service? What are other tests shall we perform on these transformers?
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09/10/2010
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Q:
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My company aims at further processing the atm. distillation residue (Mazot); and a hydrocracker unit has been chosen for this task. We need to estimate the cost of the unit and its facilities like the vacuum tower and the vis-breaker. How would you suggest we get a rough initial estimate of the costs involved?
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(5)
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07/10/2010
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Q:
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Our Sour water stripper unit is a two stage operation. The first tower operates at 7 KSCg pressure and second tower operates at 0.8 KSCg pressure. Recently we have encountered a strange problem. The color of the stripped water is milky white and also looks hazy. The overhead temperature of the second tower is running high, 100 C (Normal is 90C). Please suggest some solution.
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(2)
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07/09/2010
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Q:
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How do you calculate power density for a crude desalter unit for a given grid area? Also, is there any correlation between the droplet size and electrostatic field strength?
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(1)
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19/07/2010
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Q:
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I work in CDU with a capacity of 70000 barrels \ day I have a problem. when we introduced desalter to work, the amount of feed unit quantity decrease to 85%, especially when the introduction of washing water to work by 17000 kg \ hour DIFFERENTIAL PRESSURE on the mixing valve set to 0.35 kg \ cm2,desalter pressure is 11 kg \ cm 2 When trying to increase the unit capacity, the desalter press. Increased and send a signl to the pressure valve that is installed on the charge pump discharge to derease it’s open to reduce the increasing pressure.
Additional information:
Crude oil is taken by gravitation flow from storage tanks. The crude oil is fed pump P01 to the first part of one route heat exchanging system, where is preheated in the train of exchangers E01, 02, 03, 04 and 05 to temperature 125-130 °C. The crude oil with this temperature enters the single stage electrostatical desalting. Desalting process operates with pressure 10-12 kg/cm2g and Desalted crude is fed by means of pumps P02 to the second part of one route heat exchanging system (exchangers E06, E07, E08, E09, E10), where is heated to temperature 244-252 °C. Crude oil is dividing into four equal streams. Each of the four streams is controlled by a flow controller and enters the convection section of heater H01. In the heater (convection and radiation section) it is finally heated to temperature 342-348 °C
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(4)
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18/07/2010
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Q:
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Recently we found that the valves in the off gas (that comes from Vacuum distillation tower) separator unit line did not last long. After 6-7 months valve seat or body sprung leaks due to off gas. Please help me by suggesting the appropriate valve for this service.
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(1)
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14/07/2010
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Q:
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I work in a CDU plant. The plant loads two types of crude oil, one of paraffinic type API 48 and other one of naftenico type API 25. When crude oil naftenico is loaded high instability is had in the vacuum system (Loss Vacuum in the Column), situation that does not happen when paraffinic crude oil is loaded. Can it be something related to the ejectores? or to major production of non condensibles when crude oil naftenico loads? . The furnace-outlet temperature is 725°F with crude oil naftenico and 650°F with crude oil paraffinic.
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(1)
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08/07/2010
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Q:
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We have a very strange problem, it's that the desalter outlet crude has greater salt content than that of the inlet... the lab examinations proved that more than once...this always happens when the injection water is cut off-while switching from a tank to another. What could explain this?
Additional info/response:
1. We cut off water while switching between tanks because of the existing water accompanied with the crude from the new tank; I mean the first 30 minutes after switching to a tank, the crude has too much water to inject more.
2. How could the NaOH type could affect this situation?
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(10)
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06/07/2010
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Q:
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What is the best way of judging the efficiency of a desalter?
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(6)
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05/07/2010
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Q:
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What will be the steam dew point at 93 degC and 1.1 kg/cm2g? Pl let me know whether dew point of stripping steam used in distillation column depends upon partial pressure of steam in the total vapour mass flow going out in the distillation overhead. Pl note total vapour mass flow in our distillation column ovhd is 274000 kg/hr and stripping steam flow (at 14 kg/cm2g & 220 degC) to column is 8490kg/hr. Overhead vapour is mixture of gaseous hydrocarbon (C1 to C5 range components) & stripping steam. Pl let me know if you need more data to answer my question. Our distillation column top operates at 1.1 kg/cm2g & 93 degC.
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(3)
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05/07/2010
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Q:
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What makes outside shapes of distillation columns differ from one another? i.e. shape of pre-flash differs from CDU, CDU differs from VDU?
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(3)
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01/07/2010
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Q:
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We wish to install strainer on instrument air injection line to be used during regeneration of bi metallic pt Re reforming catalyst to avoid carry over of line material/scale on catalyst surface. Can anyone recommend what mesh size should be installed in strainer?
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(1)
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23/06/2010
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Q:
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What is the purpose of the weep hole in the chimney tray?
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(3)
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09/06/2010
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Q:
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In our atmospheric distillation unit , reduced crude recovery was constantly coming 10-16% @ 360 deg C. We increased the bottom stripping steam but we are unable to decrease beyond 10%. Are there any other ways to improve the efficiency?
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(6)
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21/05/2010
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Q:
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What are the likely effects of water carry over from desalter on Crude heater and distillation column? What steps should be taken if this happens?
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(11)
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07/04/2010
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Q:
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What is the effect of increased or decreased reflux flow on RVP of gasoline in a gasoline stabilization column? what is the effect on RVP in case column is operating on total reflux and what is the effect in case of no reflux flow?
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(3)
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06/04/2010
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Q:
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What is the solution for water carry-over from the feed tank to the desalters?
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(3)
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06/04/2010
|
Q:
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What is bone dry naphtha?
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06/04/2010
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Q:
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What is the relationship between RONC and ASTM D-86 of light naphtha boiling in the range 40C-120C? what increases its RONC? Does an increase/decrease in IBP increase the RONC or is it the increase/decrease in final boiling point that increases the RONC?
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(2)
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22/03/2010
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Q:
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What are the possible reason for column pressure (1-2 KSC pressure) fluctuation. if the column pressure increases, will the separation increase, and vice versa?
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(3)
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08/03/2010
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Q:
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In our Once Through Hydrocracker Unit, the Recycle Gas Compressor is surging from 100% opening of the anti-surge valve to 0% without any change in process parameters. It was also observed that just prior to surging the total flow at the inlet of the RGC was also increasing. We have got an amine column at the inlet of RGC suction after HP separator to reduce sulphur loading. But now due to some constraints the amine flow had to be reduced. Can anybody explain the phenomenon?
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(3)
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19/02/2010
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Q:
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I am working in DHDS. I would like to know the purpose of Carbon filter in Amine Recovery Unit. We use stripped water from Sour water stripping unit as wash water in DHDS over head coolers for dissolving ammonium salts. My query is if there are little amounts of ammonia and H2S in stripped water, and if we use the same stripped water in DHDS, will there be any problem in amine quality or will there be any effect in the quality of acid gas generated from ARU? We are facing the problem of increase in differential pressure across Carbon filter when we take stripped water in DHDS.
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(6)
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09/02/2010
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Q:
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We want to by-pass our de-salters in order to check the consequences with and without desalters on CDU. Moreover we have stopped de-emusifier dosing prior to desalters. What impacts are anticipated in your opinion and what parameters to be monitored in case when there is no desalter in crude preheat trains?
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(9)
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02/02/2010
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Q:
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What is the relationship between smoke point and ASTM D-86 of kerosene? Does an increase in IBP increase the smoke point or is it the increase in final boiling that increases the smoke point?
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(5)
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29/01/2010
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Q:
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I am engineer in a CDU. The lube vacuum column used stripping steam of 25 psig to 750F, the superheater will go to maintenance. Is it possible during the maintenance to use stripping steam of 50 psig to 750°C in this tower? This tower produce distillates for lube naphthenic and paraffinic.
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(3)
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20/12/2009
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Q:
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I'm new to the refining field. We have eletrostatic desalter with three transformers (each of 22KV). and I have questions about this:
1. Why is a single heavier transformer is not used instead of three?
2. We have 2 PSVs on desalter. What is the reason for two PSVs?
3. Why the water injection before desalter is maintained 4.5% of crude charging?
4. What is the function of grills (grids) inside the desalter?
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(3)
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14/12/2009
|
Q:
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Our furnace has 4 pass flow. Crude enters the furnace by 4" tube in the convection section. Then it changes its size by 5" X 4" reducer in the radiation section. It again changes its size outside the furnace and now this time by 8" X 5" reducer to a common header of 12" pipe line. This pipe line by a 16" X 12" reducer connected to the 16" pipe line that goes to column. My question is why we are using so many reducers in the process line?
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(3)
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29/11/2009
|
Q:
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What is the difference between the Eductor and Ejectors?
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(3)
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29/11/2009
|
Q:
|
I am a Shift supeintendent of the CDU unit.
We could not stabilize the Brine treatment package - Hydrocyclones to separate the oil and sludge from the Brine of the Desalter outlet.
If anybody have the experience regarding the operations of Hydrocyclones (Brine treatment package) in the Brine system, please share with me.
If I can get the optimum dela pressure across, it will be helpful; I could not follow the vendor operational guidelines as it is not performing good.
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(3)
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16/11/2009
|
Q:
|
For FCCU Fractionator bottoms to slurry circulating pumps, is there a preference of side inlet/bottom outlet coke strainers over side inlet/side outlet coke strainers?
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15/11/2009
|
Q:
|
For the FCCU Fractionator Column bottoms to the slurry circulating pumps, what are the preferred coke strainer types between side inlet/ bottom outlet coke strainer and side inlet /side outlet coke strainer? Are there any precautions to be taken for use of side inlet/side outlet coke strainers?
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05/10/2009
|
Q:
|
I'm researching about paraffin wax oxidation. We're experiencing a phenomenon of lost of colour in the hydrotreated waxes.
What kind of oxidation phenomena can I have in a hydrotreated paraffin wax?
Has anyone experienced this problem with paraffin waxes?
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(2)
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13/09/2009
|
Q:
|
What is the effect of phenols in desalting operation? We have a stripped sour water having a phenol content of 600 ppm which is to be used as wash water in the upstream of desalter.
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(4)
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11/08/2009
|
Q:
|
My question concerns narrow or "light" naphtha. As a broker and trader, most of the product I see has an IBP (initial boiling point) low range of 40 degrees Celsius. I have a client seeking to purchase product with specs stating 35 degrees. I believe this to be highly unusual, or is this a common specification? Please advise.
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(2)
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31/07/2009
|
Q:
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I am doing some research on the Cameron acquisition of Natco and was interested in learning if desalters are sold in the United States. My understanding is that they are installed in refineries when the refinery is built and inasmuch as there is no new refinery construction there are no desalter sales in the USA. Is that correct?
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(3)
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07/07/2009
|
Q:
|
How is commercial hexane i.e. hexane with n-hexane content of 40 to 60 %, produced ?
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|
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03/07/2009
|
Q:
|
One of the observations pertaining to 2 stage desalters is that 2nd stage efficiency remains poor as compare to 1st stage (50 to 60 % in 2nd stage as compare to 91 % in 1st stage) despite trying all necessary parameter adjustments such as Delta P, Fresh water flow, interface level, emulsion checking etc.
Is there a means of remedying this discrepancy?
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(3)
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24/05/2009
|
Q:
|
In order to achieve higher gas oil recovery, we are working on a conceptual project of installing a vacuum flasher on the D/S of VDU. VR from VDU column bottom will be heated and flashed in a separate column having flsha zone pressure lower and temperature higher than that of VDU column flash zone. Preliminary simulations suggests the technical feasibility. We want to confirm that whether such a set-up has been installed anywhere and what may be the foreseen challenges?
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(2)
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15/04/2009
|
Q:
|
This question is about heated and heatless instrument air dehydration packages.
While heated dehydration systems rely on blower/heater combination for regeneration, heatless systems require a dry instrument air stream for regeneration (up to 15%) which leads to a larger compressor to ensure a steady supply of IA.
Which of these heated/heatless systems is better and why? Are there significant lifecycle cost and availability/reliability issues to differentiate?
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19/03/2009
|
Q:
|
With some experts projecting crude prices to creep back up to $75/bbl by mid-summer 2009, should we expect to see a higher level of refinery intermediates (e.g., heavy gas oil, "lifted" DAO, etc.) being exchanged among "networked" refining facilities?
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12/03/2009
|
Q:
|
We are experiencing excessive backwash frequency on our auto backwash filters (25 micron size) on the feed to the diesel hydrotreater unit.The hot feed is a blend of straight run kero, LDO and HDO which is fed directly from the crude unit with no makeup from intermediate storage. The feed when analyzed indicates a particulate level of about 6 ppm which in my opinion is low to cause such a problem. Has anyone experienced similar phenomena when the moisture levels in the feed are high? Moisture when analyzed was observed to be about 750 ppm in the feed.
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(3)
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19/02/2009
|
Q:
|
We have two sets of desalters and each set has 2 stages i.e., 2 desalters in series. Our crude oil props input to desalter mainly BS&W, API, Salt Content, Filterable solids, Viscosity and Residence time are within (or equal to, in case of R.Time) the design conditions. But we are unable to obtain the desired outlet parameters even though these are within the design parameters. For Ex: Salt ptb (o/l) - Design:1, Desired: <0.4, but actual is 0.6-0.8, like that there are variations in all parameters viz., BS&W, Filterable solids & Oil in water etc.
How can we improve these parameters in the existing train so that we can reduce the severe pre heat train fouling that is very frequently happening and overhead corrosion. Would effective demulsifying and anti foulant agents will be helpful in this regard? If so, please recommend?
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(2)
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12/02/2009
|
Q:
|
A coalescer is required to be installed in straight run kero line of the CDU. What are the requirements under OISD (Oil Industry Safety Directive ) for this? Are there any other safety directive norms that need to be considered?
|
(1)
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07/02/2009
|
Q:
|
Where can I obtain information about Vacuum distillation unit overhead sourgas minimization?
What are the parameters that effect the sour gas generation rate? Are there any correlations available to relate those parameters to sourgas rate?
What are the methods and ways to minimize the cracking of reduced crude oil in vacuum unit charge heater? what are the main effecting parameters of fouling the vacuum charge heater?
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(4)
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06/02/2009
|
Q:
|
What are the best preventative measures for avoiding coking in wash bed section of vacuum tower and corrosion of top section of CDU?
|
(3)
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|
18/12/2008
|
Q:
|
What is the best distributor for water feed to a atmospheric condensate drum to avoid steam losses to ambient?
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|
|
23/09/2008
|
Q:
|
I am working a project where I am trying detect phase changes. The project consist of detecting phase changes from water to butane by using flow meter density detectors. This idea is only for ideal case, but the reality is that, caustic may be present. Here is where the issue comes.
The question that I have is this: what method should I use to detect different phases. For example, mixed water and caustic? mixed Butane and Caustic? Again, the point is to detect phase density changes from water to butane.
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08/08/2008
|
Q:
|
Are recent improvements to FCC cyclone technology adequate enough to lower solids concentration in slurry oil down to a level (e.g., < 250 ppm) that the slurry oil can be sold without need for filtering?
|
(1)
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04/08/2008
|
Q:
|
How effective have membrane separation systems integrated into recent clean fuel strategies been in reducing sulphur levels, octane upgrading, etc.?
|
(1)
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|
07/07/2008
|
Q:
|
In general, where has the influence of good fractionation allowed for significant improvements in meeting stringent petrochemical product specifications (e.g., propylene, styrene, etc.) at higher charge rates? Besides the recent improvements to fractionation column internals, what is the extent to which automation & control systems can be leveraged to deliver higher efficiency, run-lengths and resistance to corrosion in product recovery trains?
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|
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22/06/2008
|
Q:
|
We have a crude preflash column where crude after being heated is flashed.
the column has three streams. One is the top which is light naphtha which is taken from the reflux drum via reflux pump as one stream. The other is sent as reflux on temp control The column has a fired reboiler.
A side cut is Heavy Naphtha below Hnaphtha cut is a pump around. Bottom is kerosene plus which is separated in another tower.
We wanted to reduce the EP of lt naphtha. We carried simulation on hysis and were getting the desired EP using HEPT OF 2 FT for CMR 1 as there is packed bed of 10 ft between LT and H naphtha we were getting five theoretical trays
on the plant we adjusted all perimeter as per simulation but we could not get even close to it the ep remained high. Increasing reflux severalfold could not achieve the end point.
We took delta p across the bed. It is low and pct flood predicted by Hysis is 16pct far from flood.
We reduced capacity but no avail. Could it be low flood which is responsible? We want to check all angles before we open it up. There are no gamma scan facilities available so we can't do a scan.
Can someone suggest what angle to look for?
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(1)
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21/06/2008
|
Q:
|
Is there a noticeable increase in blending clarified FCC slurry oil into No. 6 fuel oil? Since this obviously circumvents the need for blending lighter, higher-value products into the No. 6 fuel oil, how much of an impact on total refinery profitability can be expected? Are some refiners instead opting to use higher percentages of slurry oil as feedstock to a coker unit or a hydrocracker?
|
(1)
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28/05/2008
|
Q:
|
Can someone provide guidelines on the design of a vacuum column for light waxy atmos bottoms? Will it be any different than designing a vacuum column for heavy crude oils with high metals etc> The atmos bottom that we are considering is waxy with no metals and very low sulphur but we have to limit wax in the lvgo so that it is used as diesel.
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|
05/02/2008
|
Q:
|
Heavy crude oil desalting in electrostatic desalter designed for normal crude creates interface level problem and results in more oil in desalter effluent. What best operating and design practices should be followed to overcome this problem?
|
(8)
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21/01/2008
|
Q:
|
How can I predict HETP for Sulzer's structured packings (BX, Mellapak 250.Y or EX) when reflux is not total, i.e., when some distillate is taken off (e.g., 10, 20, 50 or 75 %)? Does it depend on the mixture to distill or is it an inherent characteristic of the packing?
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(2)
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19/01/2008
|
Q:
|
We have a topping plant with capacity of 40,000B/D. One Desalter is installed before the Flash tower with emulsion breaker injection on the inlet pipe for desalter.
We are looking for the suppliers for overhead corrosion control chemicals (NALCO filmer & Neutralizer) for the flash tower and the emulsion breaker
|
(1)
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|
08/01/2008
|
Q:
|
Could the use of a 70 micron Element Filter (Wedge Wire), in place of a 25 micron Filter, for the filtration of hydrocracker unit feed, result in any problems?
|
(4)
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26/11/2007
|
Q:
|
During the percolation process, the tower is packed with bauxite, then when saturated, cleaned with naphtha, then unpacked and the bauxite burned in a rotary oven at 550 Centigrades. Then the bauxite is re-activated with sulphuric acid and the tower packed again. Is this correct procedure?
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(1)
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09/11/2007
|
Q:
|
Please tell me the process for handling slop oil. How much is produced per barrel, what is the process to separate it back to crude, disposal, economics of this issue? What are the problems oil refineries face with this issue?
|
(1)
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|
04/11/2007
|
Q:
|
Neutralizing amines are used in PH control in crude distillation over head systems to minimize corrosion. While studying the possibility of using such amines for Vacuum distillation several contradictory points of view appear some approving and some objecting. What do you suggest? We currently use ammonia in our vacuum distillation.
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(2)
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|
25/10/2007
|
Q:
|
What systems are available for H2S removal from crude oil (39 API gravity crude measuring 50 ppm H2S at the wellhead)? Where can I obtain quotes for pricing the necessary equipment?
|
|
|
15/09/2007
|
Q:
|
What processes are available for
1. the separation of oil from slack wax
2. the separation of wax fom residue wax
3. the hydrogenation of wax?
|
(1)
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|
07/09/2007
|
Q:
|
What technologies are available to reduce the VOCs from the truck loading? Our truck loading is bottom loading type.
|
(2)
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|
06/09/2007
|
Q:
|
In the off gases from our vacuum distillation column hydrogen % has been up to 30-35% by volume.This vacuum unit is mild severity dry distillation with designed VGO end point of 510 deg C.
The overhead boot water PH also remains on the lower side (~5) even though the neutraliser is added in large quantities (more than 100 ppm). The same neutraliser has used earlier for the same type of crudes.
Has anyone had this type of experience? What may be the reason for the same?
|
(1)
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|
06/08/2007
|
Q:
|
In certain gas processing installations, we find that the Pressure Safety Valves (PSVs) on demethaniser, deethaniser and ethylene towers vent directly to the atmosphere. Is this acceptable practice or should PSVs always be connected to flare systems? What is best practice for routing of safety valve discharges of such columns handling lighter hydrocarbons?
|
(6)
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|
06/08/2007
|
Q:
|
In certain gas processing installations, we find that the Pressure Safety Valves (PSVs) on demethaniser, deethaniser and ethylene towers vent directly to the atmosphere. Is this acceptable practice or should PSVs always be connected to flare systems? What is best practice for routing of safety valve discharges of such columns handling lighter hydrocarbons?
|
(6)
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|
30/07/2007
|
Q:
|
Fusel oil is generated in Crude Methanol purification/distillation. It has about 50% water and balance is various alcohols. To use it as a burner fuel is very taxing as the mass flow rate is very high compared to the fuel heat value.
The question is - How can this mixture be concentrated in an economic manner such that the water component is removed? This concentrated Fusel oil shall have much higher heat value and it shall improve burner firing efficiency.
|
(2)
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|
23/07/2007
|
Q:
|
How can the capacity and efficiency of an existing deisobutaniser be increased?
|
(3)
|