Q & A > Gas Processing and Treatment
Date  Replies
08/06/2017 Q: In case of PSA expansion with one pair, is it a must to have the same PSA skid between the new vessels in size (length)?  
21/04/2017 Q: In an atmospheric CCR reforming unit how can measure the chloride content (HCl) of the regeneration vent gas?
If you have a Chlorsorb system how can you measure its efficiency?
(3)
08/03/2017 Q: There are 6 identical trains processing natural gas bringing H2S < 4PPM and CO2 From 5% to 2%. All parameters are same. MDEA is the solvent used. But interestingly out of the 6 absorbers one is not properly functioning in terms of capacity and quality.The lean amine in all other absorbers is pale yellow and the one having problems is brownish. The solvent tested for all parameters, found ok except colour. Could any one come out with a solution? (2)
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? (2)
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. (2)
22/09/2016 Q: We have problem of failing silver strip corrosion in storage tank of jet fuel and same time rundown sample is always clear. Please suggest solution. (3)
28/07/2016 Q: I am working as production Engineer in FCC block which consists of FCC and LPG Treater unit.
LPG treater treats cracked LPG from FCC and Delayed coker Unit.IT has capacity of treating 404 m3/hr.We are experiencing a peculiar problem these days. Our LPG is getting contaminated with light yellow colour at the outlet of LPG treater unit. LPG ex FCC and coker unit is clear and having no signs of yellowish colour but at outlet it is failing. Copper strip corrosion and weathering tests are absolutely fine.
The following are the observations --
1) These units are commissioned 6 months ago and we didn't come across DSO formation in DSO settler but LPG quality is ok (Probably due to low mercaptans load in LPG, DSO is not forming in the process)
2) The lean amine H2s loading in FCC & DCU LPG washing is around 1200 -2400 ppm. May be noted here that on earlier occasions with this loading of amine,the downstream treater unit was giving good LPG wrt all lab tests.
3)The only problem is with colouring with which we are struggling.
4) sometimes we experience yellowish colour at the unit inlet of LPG treater.Even at the outlet of FCC & DCU also.But the colour is not that much remarkable.But in treater outlet the pale yellow/dark yellow colour is visible upon complete evaporation of LPG.Sometimes 0.5 to 0.1 ml of yellow residue is left and sometimes not.
5) We took complete shutdown of unit, depressurised, all caustic drained from system and vessels and fibre film contactors were water washed and again took the system inline but the yellowish colour in the system LPG still coming.The custic regenration is good and quality is good and copper corrosion test is ok.
The reason for yellowishness we unable to find out.
Can anyone share their experiences in dealing with this type of typical problem in MERICHEM section of caustic wash of LPG?
(4)
08/06/2016 Q: A trip is provided on high tail gas temperature of Sulfur Recovery Unit. It will bypass the amine system & tail gases will be directly routed to Incinerator. Why this trip is provided?
(2)
11/12/2015 Q: We are facing frequent choking problem in phenolic water stripper bottom pump strainer. Phenolic stripper is operated at 1.4 kg/cm2 of top pressure. Column top temp is 93 deg C & bottom temp 129 deg C. Column bottom section has reboiler which use saturated LP steam (157 deg C). It is prudent to note that phenolic water feed pump is not getting choked. Can anyone have a solution to overcome the choking issue?? (2)
22/11/2015 Q: We have some LPG Merox units with amine absorber before the Merox unit. We use MDEA in the amine absorber and we have experienced some problems of amine carryover in the LPG.
Can anyone comment on the impact of the amine contamination in the Merox units? Besides the possible formation of emulsions, could there be any other problem?
(6)
13/11/2015 Q: We are suffering from water carryover with tail gases form quench top in tail gas treater unit of Sulphur Recovery Block.
In case of absorber bypass due to S/D of Common Regeneration Unit; tail gases are routed from quench tower top to incinerator which results in water accumulation inside incinerator which is a major problem.
To tackle this it has been planned not to bypass absorber and flow tail gases from absorber without any amine flow.
I wish to know what could be possible ill effect or problems with the same?
(1)
04/11/2015 Q: What is the test method for determining hydrocarbons in amine solution? (1)
21/10/2015 Q: What is the acceptable level of hydrocarbon removal (%) across the Activated carbon filter in Amine slip stream?  
21/10/2015 Q: What is the maximum level of hydrocarbons permissible in lean amine so as not to cause foaming in the FCC offgas Amine absorber? (2)
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)
07/09/2015 Q: What is the difference between weathering and RVP? (3)
28/08/2015 Q: We have Refinery off gas PSA processing CLPS off gases from hydrocracker and CCR off gas at 23 kg/cm2g. Tail gas from this PSA is at 0.5 kg/cm2g which is compressed to 4 kg/cm2g using oil flooded screw compressor and routed to fuel gas header.
During upset or flow variation in PSA frequent LO filter choking is observed for the screw compressor. Sometimes the filter choking is so fast that compressor trips on high differential pressure between LO and gas. What could be the reason for such high filter choking during particular time of PSA variation. We then clean the oil using continuous filtration and it takes 2-3 days to again normalize the things. Once the lube oil DP gets normalized it continues to run longer without any LO filter change out. Blackish deposits are observed on filter cartridges. Is anyone facing similar issue in tail gas screw compressor? What remedial actions are available to resolve the issue?
(2)
25/08/2015 Q: We have safety valves on LNG discharges tank discharging to atmosphere through a tail pipe approx. 25m long.To prevent ignition of the release in the event of a lightning strike, would it be acceptable to provide a flame arrestor on the vent? (3)
16/06/2015 Q: We are observing high CS2 content in our straight run naphtha. This is not on regular basis but frequent and sometime it goes up to more than 20-25 ppm also. Please advise what can be source of such high CS2 content in naphtha intermittently.
The sources may be narrowed down to:
1. Presence of CS2 in Crude itself- Please suggest the probability of the same and if any known crude with high CS2?
2. Since CS2 formation requires very high temp, can it be formed in crude heaters? or any other process?
3. Though probability is less, can it come from recycle hydrotreated naphtha?
(1)
05/06/2015 Q: We have floating roof naphtha tanks. I want to calculate floating roof weight displacement volume with, density at 15 degc, weight of floating roof. Please suggest a formula.  
06/05/2015 Q: How do I check whether the water cooler tubes have punctured heat exchanger? Cooling tower is tapped with various units like GSU, DPD, SRU etc. Gas is coming in cooling water return line. The unit in-charge says our units are ok. Can we check at heat exchanger? How?
(4)
29/04/2015 Q: We are facing issue of LPG product offspec due to total sulfur high. we have a UOP LPG merox treamtent unit (LPG enter from CDU to absorber> filter> coalescer> Prewash>Extractor>sand filter).
design RSH in LPG feed to Merox is 800 wtppm. but currently due to feed, we are processing upto 5000 wtppm. which parameters to be checked thoroughly in order to reduce this offspec.
1. Regen caustic quality (Mercaptide & Disulfide) feed to extractor.
2. LPG to Prewash ?
3. Lead acetate test of which streams will help us to identify at field rather than sending samples and wait for results.
(4)
23/04/2015 Q: We process condensate to produce ATF. For 2 months ATF color is deteriorating day to day. What could be the reason? (2)
13/02/2015 Q: In the CO2 recovery section of Ammonia plants, using MDEA, are there any advantages to keeping the exchangers using cooling water on the tube side (e.g.Lean solution cooler, LP flash gas cooler) vertical, instead of horizontal? (1)
13/02/2015 Q: In the CO2 recovery process using MDEA solution , in ammonia plants, is it always the practice to mount the LP flash drum directly on top of the HP flash drum, instead of separate mounting? (1)
21/01/2015 Q: We have a kettle type reboiler with weir plate and its liquid outlet is going regenerator in benfield process of co2 removal in ammonia plant. Is it necessary or required to put vortex breaker in the reboiler liquid outlet? Which design is applicable for designod reboiler? (1)
30/12/2014 Q: This question is related to kerosene merox unit. After processing kerosene in merox unit, what are the main reasons for poor saybolt color of kerosene product? If kerosene feed to the merox unit has saybolt color of +26, kerosene product from merox unit observes saybolt color of <16. Can someone explain the possible compounds which causes color problems to the kerosene product? If we go to Kerosene Hydrotreater, there will not be any issues of color problems and in fact it will be improved because of olefin and aromatic saturation. Please share any literature or chemistry related to the kerosene color problems in merox units. (3)
24/10/2014 Q: What is Nelson Complexity Index factor for ATF,Gasoline and LPG Merox , Sulphur Recovery Unit process ? (1)
24/10/2014 Q: What is Nelson Complexity Index factor for ATF, Gasoline and LPG Merox , Sulphur Recovery Unit process ?  
25/09/2014 Q: Can we process FCC's Clarified oil (CLO) or Decant oil as feed to Hydro cracker? My question is that Unconverted oil from Hydro cracker is usually good feed to FCC, So I would like to know if we process FCC CLO in hydrocracker then how much of it will it to convert to Unconverted oil in Hydrocracker? We will use filters to reduce catalyst content in CLO so that hydro treater won't get affected. (2)
13/08/2014 Q: 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?
(2)
03/07/2014 Q: Does anybody use MDEA on Amine treating on FCC?
We proceed hydrotreating feed on FCC, and we use DEA on Amine treating. We want to switch DEA with MDEA.
What should I pay attention to during this switching?
(2)
13/04/2014 Q: We have a Steam methane reformer having side fired self respiratory burners. To attain the correct O2 in flue gas of primary reformer, burner dampers are being adjusted. What is the correct sequence for throttling the burners? Should the bottom most burners should be throttled more than the top ones or vice versa?  
11/04/2014 Q: In case of side fired self respiratory burners in reformers what is the correct sequence of adjusting the air?
From bottom row burners to top row burners in increasing trends:
in 1st row 40%, 2nd row 40%, 3rd row 40%, 4th row 30%, 5th row 30% & 6th row 30%
OR
in 1st row 30%, 2nd row 30%, 3rd row 30%, 4th row 40%, 5th row 40% & 6th row 40%.
This flue gas is going to convection section for heat recovery.
 
29/01/2014 Q: We are having 3 nos. of identical SRU trains (2 nos. of Claus & 2 nos. of CBA reactors) with each having a capacity of 65T/day of Sulphur production. We are frequently facing a problem of second condenser tubes leaks problem and plugging of rundowns due to catalyst dust carry over. Operating temperatures are being maintained above dew point in all three trains. However, tubes leaks problems are coming in only one train very frequently. Can anyone suggest the cause for frequent failure of condenser tubes leaks?
C2 outlet temperature is being maintained >150 degC against the design of 168 degC. 1st claus reactor (R1) outlet temperature was being maintained >340 degC. C2 outlet temperatures observed >150 degC even for R1 o/l temperature of 310-320 degC.
In spite of maintaining C2 o/l temperature >150 degC, sudden decrease in C2 o/l temperature and increase in system back pressure is being observed due to tubes leaks.
(3)
13/01/2014 Q: We have steam methane reformer. The outer surface of the tube is having deposits and leading to high fuel consumption as well as high temperature in flue gas side in waste heat section. During turn around we want to clean the outer surface of reformer catalyst tube so that we can reduce the fuel consumption and reduction in waste heat section temperatures. Is there any standard method available to clean the outer surface of the tubes? (5)
12/11/2013 Q: In DHDT unit, RGC anti-surge control valve is opened at around 13-15 %, at this opening deviation from surge line is 0.14-0.20. The design molecular weight of Recycle gas considered is 2.94 whereas actual Recycle gas molecular weight is in the range of 2.25-2.4. Suction temp/ Pres:63 deg C /110 kg/cm2;Discharge temp/ Pres:90 deg C /131 kg/cm2
1.0 Can we fully close the anti-surge valve in order to increase energy efficiency of RGC ?
2.0 what other actions can be taken to minimize RGC anti-surge opening ?
3.0 By Incorporating the actual recycle gas molecular weight in anti surge controller block and compression suction flow transmitter, will there be any improvement in deviation from surge line ?
(2)
01/10/2013 Q: I need a cost estimation for hydro desulfurization unit (HDS) by amine (DEA) in a petroleum refinery
Given data for this unit :
-hydrogen sulfide in feed = 8.6 mol % = 86177 ppm
- feed =890 T/D
-target.< 200 ppm
 
01/10/2013 Q: I need to know the overall cost for the hydrogen sulfide removal unit by DEA,...And individual cost for each unit  
19/08/2013 Q: We are heating desulphurization unit by natural gas (NG) without hydrogen up to 300deg C and being vented in to the flare. Instead of flaring this NG, after coming out out from desulphurization unit can we cool and compress in NG compressor and heat in waste heat section of reformer and feed into desulphuriser unit for heating? Please mention advantages and dis adventages? one of the advantage is vent of NG can be stopped. Any effect on catalyst? Like coke formation etc.

Further info:
We are only using only NG (CH4-93%+ 7% N2) first in NG compressor and then heating into waste heat section to increase the temperature of NG for heating desulphurisation section. After the heating it is being vent into flare. My opinion is instead of venting into flare can we cool it and again in compress and heat in waste heat section and again feed to desulphurization section to increase the temperature up to 300deg C. Any adverse effect on catalyst etc.? With this arrangement we can avoid venting of precious NG in to flare. No hydrogen is added during heating process.
(2)
19/08/2013 Q: We are heating the desulphurisation unit with NG without hydrogen up to 300 deg. C and then NG is being vented through the flare. We want recycle this NG by cooling into exchanger and compress in to NG compressor again and heat in reformer waste heat section and to desulphuriser unit. Is this method ok?  
31/07/2013 Q: Would using a molecular sieve be the best bet for removal of chloride salt contaminants of refinery fuel gas? We are consistently plugging fuel gas valves, strainers, and burners causing reliability issues with our fired heaters. (1)
30/07/2013 Q: We are currently experiencing continued plugging of our refinery fuel gas control valves, strainers and burners. We went through a re-org 3 years ago where one of reformers was brought down. Since then, we have seen an increase of chloride salt contaminants to our fuel system from our other reformer. We currently run 2 molecular sieves in series on or hydrogen header. I proposed to increase heater reliability and reduce chloride salt contaminants to take second mole sieve and pipe the fuel gas header to it. This would of course be after testing hydrogen chloride content with only one sieve in service and projecting those results on compressor reliability from our maintenance group. If no real future damage can be projected and current single phase mole sieve can handle hydrogen system, would a mole sieve for the fuel gas be an adequate route since the vessel is already there and would only require a piping mod?  
02/07/2013 Q: What are the methods and guidelines to predict SRU Claus Catalyst life. (2)
06/06/2013 Q: What is the effect of high moisture (free water) on naphtha hydrotreater catalyst (Ni-Mo) performance?  
12/05/2013 Q: My question is on Acetylene Selective Hydrogenation Catalyst (Palladium –Pd based with promoters):
Ethane gas gets cracked in the Cracking Furnaces and the effluent goes through series of processes that includes quenching, heavy contaminants / heavy hydrocarbons removals, Multi-stage Compression, Caustic Scrubbing with Drying leading to De-Ethaniser (DeC2), and DeC2 Column Overhead vapour to the Two-stage Acetylene Hydrogenation Reactors. Main feed Ethane gas has a spec. of CO2: 200 to 1000 ppm; Total Sulfur: 500 ppm; Moisture content: 100ppm and it is directly cracked in the Furnaces. There are other feed streams having Sulfur ppm in the range upto 50 or so, with metal traces at lower ppb levels. The Reactors are operated with Carbon Monoxide level of 1000 ppm to 3000 ppm Max or so, at the upset conditions. Outlet Acetylene ppm levels are stringent in the range of 0.2 to 0.3 to produce Ethylene with 1 ppm Max Acetylene impurity.
a) Pl. let me know what all process parameters have direct impact on Catalyst deactivation and thereby short run-time requiring ex-situ Regeneration.
b) How will you control the parameters effectively to have much longer Catalyst run-time?
c) What is normal catalyst run-time for such Catalysts irrespective of any Catalyst vendors?
d) Whether going for Regeneration, would it be recommended to revive activity and selectivity to that of fresh material? Any risk involved in taking decision in favour of Regeneration?
e) Vendors confuse often with jargons, Reactivation and Regeneration. Are they one and the same or the process of reviving the spent material to the active phase to prolong the operation with recycle not only due to downtime of plant but also, expensive nature of catalyst with precious metals?
f) Pl. suggest suitable catalyst vendors with whom development activity can be collaborated with the company’s R&D Centre.
g) Any other important points in relation to specific Catalyst poisons, improving run-time atleast upto 4-5 years if not 10 years+

Your thoughts on this, in whole or part, greatly appreciated.
 
19/04/2013 Q: Why do multistage reciprocating compressors have different compression ratio for each stage? (1)
19/04/2013 Q: In the case of multistage compressor, why is it that the compression ratio are not set equally?  
12/03/2013 Q: In one of our FCCUs we have problems closing heat balance due to the processing of a very hydrotreated feedstock. We have to use torch oil (LCO or fresh feed) to maintain regenerator at its minimum temperature.
We are evaluating the possibility of using other feedstock as torch oil. Has anyone experience in using fuel gas or natural gas as torch oil in the regenerator? What major modifications in hardware are required?
(2)
24/01/2013 Q: Recycle gas from high pressure separator is being treated in amine absorber. Treated gas then routed to recycle gas compressor (2 stage). We are observing lot of liquid accumulation in the inter stage KOD (after cooler). Liquid sample is analyzed and found that 99.7 vol% of sample contains water and remaining as amine. We are adding water to the amine system to compensate the loss through inter stage KOD.
Lean amine to absorber temperature is being varied in the range of 55 to 60 °C to maintain design DT between amine & gas. We have observed that the absorber top temperature is high when compared to bottom temperature.
Why only water is getting carried over (in huge quantities) along with the treated gas? Is the same observed in any high pressure amine absorbers.
Why temperature profile in the amine absorber is in the opposite direction? Is this related to moving of heat of absorption profile from bottom to top?

Further info:
If there is a foaming in the column the following should have happened:
1. Fluctuations in absorber bottom level; not fluctuating
2. Severe fluctuations in DP across the absorber; not fluctuating
3. Improper stripping of H2S from gas; not observed
4. Amine foaming tendency test is also done and found satisfactory.
In the case of foaming in the absorber, amine should also get carried over along with water. But this is not happening. As i said earlier amine content in the water sample is in the range of 0.3 to 0.5 wt% only.
What is the ideal delta temperature to be maintained between gas and amine?

(2)
09/01/2013 Q: In a Ammonia storage tank (at atmospheric condition and -33 deg C) when will more boil-off happen:
a) If the Ammonia is filled till half of the level , or
b) If Ammonia is filled up to full height
Tank construction is with Double wall, with perlite concrete at bottom, foam glass at bottom and Mineral wool at suspended deck.

(1)
21/10/2012 Q: We encounter polymerisation from phenolic material in the KO drum of the SWS overhead system. As a result, KO drum bottom is frequently plugged and this means that in case of massive carryover from the SWS reflux drum the liquid would not be pumped out of the system. Source of phenols is the FCC. The analyses of the polymeric fouling don't give clarity on the other components, but the appearance is similar to phenol-formaldehyde resins.
Please note SWS reflux drum is operated at 87 C (to avoid ammonium bisulfide deposition) and the downstream system is traced and heated with steam so that the temperature is normally > 100 C (no water is separated in the KO drum in normal operation).
Did anyone encounter similar problems in SWS OVHD systems, and how did you solve them?
(3)
25/09/2012 Q: Why is nitrogen blanketing provided for the oxygen-enriched line in a sulphur recovery unit?  
17/07/2012 Q: What is the main difference between Accumulation and Over pressure for a relief valve? could anyone explain their importance while sizing a relief valve. (1)
17/07/2012 Q: What is the significance of PSV's discharge coefficient? how it will impact on relief valve sizing?  
17/07/2012 Q: Could anyone explain the significance of PSV's %blowdown value that we need to mention when sizing a Pressure Relief Valve.  
31/05/2012 Q: Our sulphur plant is two converter, three condenser designed for producing 180 mtpd sulphur with recovery of 96% and downstream TGTU for tail gas recovery. SRU have very frequent choking in condenser III. Deposits are very hard in nature and can not be removed even with hydrojet. This have made the condenser suseptible for leakages.
We operate SRU with acid gas from ARU. Amine is used in HCU and FG, LPG aboserbers treating LPG and FG from DCU and MSB, CDU. Solution is MDEA with concentration of 19%. Oil is also observed in amine and filtration system is not very effective. Amine could be source for initiation of problem in MCC but not able to understand the salt which is depositing in the condenser III. What could be the reason of salt deposits in condenser III?
(2)
04/05/2012 Q: Post Turn Around in Mar-2010, we have been facing issue of C3/C4 splitter re-boiler fouling. This reboiler is at down stream of LPG merox unit. Simultaneously, high delta P has been observed across LPG sand filter. Recently we have opened up the LP sand filter and found some sticky material showing some of fouling material is being slipped through sand filter (partially by passed at times) and accumulating in reboiler. What could be the reasons of fouling in LPG Merox?
Analysis of the sludge collected from the re-boiler was showing mainly HC, 1.5% sodium, ppm levels of silica and iron.
More over lead acetate test was carried out at outlet of sand filter and it was negative showing no sulphur being slipped in treated LPG.
(2)
13/04/2012 Q: How does the MCC (metal-catalysed coking) phenomenon happen in a high temperature fired heater in CCR unit, and how to protect? (1)
12/04/2012 Q: Please explain to me what the MCC (metal catalyzed coking) in a high temperature equipment of a Naphtha Platforming unit is, and how to prevent it? (5)
28/01/2012 Q: In one of our FCC units (Kellog Orthoflow model), we are suffering severe problems of fouling (fines deposition) in the turboexpander. The scheme of the flue gas circuit is: two stage cyclones in the regenerator + Shell Third Stage Separator before turboexpander + 4th Stage Separator (cyclon) to recover flue gas from fines coming from TSS.
We have also observed high level of moisture in the fines from 4th Stage Separator (10-15%wt). So we suspect that the fouling of the expander is due a cold point in the flue gas circuit (where flue gas humidity is condensed) or an uncontrolled inlet of water / steam.
Has anyone experienced this kind of problems in an FCCU? What could be the potential causes of the severe fouling of the expander?
(2)
13/01/2012 Q: Can DIPA will be used for TGTU applications? Who can do basic design of units with DIPA? (1)
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.  
21/12/2011 Q: My question is related with high sulphur content in LPG from crude distillation.
In one of our refineries we have detected high sulphur content in LPG from crude distillation. The scheme is as follows: Crude distillation - Gas concentration unit - Debutanizer - Amine absorber - Merox extractive. The high S content is mainly due to high dimethylsulphide (DMS) and dimethyldisulphide (DMDS). We measure high DMS and DMDS both at the entry and outlet of the amine absorber and LPG Merox. We have also seen some unexpected behaviour with these species: DMDS increase through the amine plant (DMDS higher in the outlet than in the inlet) and decrease in the Merox unit. The same with DMS. But sometimes we have also observe that DMDS decrease in the amine plant (??)
My questions are:
- What could be the origin of DMS and DMDS (synthetic / heavy crudes, slops processing...)?.
DMDS could be re-entry sulphur in Merox, but we have observe it in the inlet of amines and Merox (It seems that both compounds come with the crude)
- Could be DMDS come from oxidation of methylmercaptan in the topping, Gascon or amines (where there is not Merox catalyst) if oxygen is present in the LPG?
- If DMDS is in the crude, according to its boiling point it should end in the heavy light / heavy naphtha. Has anyone observe high DMDS in LPG in his refinery?
- Could DMS and DMDS increase or decrease in the amine plant or Merox? (I do not think so). Are these compunds partially soluble in NaOH or could be removed in the sand filter?
- DMDS could also come from the circulating NaOH in Merox plant, if quality is not good (high concentratrion of disulphide in NaOH)? What is the normal or recomended concentration of disulphides in regenerated NaOH?
- What could be the alternatives to remove these compounds? I expect that nothing can be done in amine / Merox, these compounds are not reactive.
(3)
24/11/2011 Q: Currently, I am working on proposal of Ammonia flare system project for one of the client based in GCC. I have an experience on the HC flare system.
There are 2 flare header coming to the ammonia flare system as below.
1. PSV discharge having liquid + gas stream
2. PSV discharge having only Gas stream
The PSV discharge having liquid + gaseous stream will be transferred to the Flash tank for separation of liquid & gas stream. The separated liquid stream will be sent to ammonia storage tank & Gas stream connect with the main gas flare header & transferred to Ammonia KOD.
I have following question regarding the system, may u consider silly also
1.Why water seal drum is not considered to in this design? HC flare system have a separate KOD & water seal drum for adequate protection against flash back form the flare tip while Ammonia have only KOD integral with flare stack.
2.What is reason for providing the steam tracing to the flash tank?
(1)
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
(1)
16/09/2011 Q: What is the best way to treat waste water that contains high sulfinol to discharge level? (2)
16/09/2011 Q: I have a question about DeSOx Unit of RFCC
Our plant has a DeSOx unit that removes SOx and spent Cat’ (=dust) in Regenerator flue gas to meet environment standard.
After removing SOx and spent Cat’ by Mg(OH)2 solution, the waste water that includes suspended solid like spent cat’ is removed through filter press.
Filter press is dual type. When one is working, the other is stand-by. (Running time :18~30hr)
Because operation time of the two filter presses is not fixed and unknown, the cleaning man has to stay on or near the filter press to clean it, when switching.
So I want to ask:
1. How do you treat resulting waste water to meet environment regulation?
2. If you use filter press, what is the best way of managing it?
(3)
05/09/2011 Q: Can anyone give me a tip off on how to reduce high COD between 200,000-500,000 from sulfenol/amine waste in the gas plant.
(3)
27/06/2011 Q: Can somebody tell me what is the shelf life of carbon disulphide CS2 chemical?  
20/06/2011 Q: We need to revamp our NHT. Before revamp: 23500 bpd SR Naphtha 100ppm S, Naphtha product for CCR feed has 0.5 wt ppm. After revamp: 30 000 bpd (90% vol SR Naphtha, 10 % coker Naphtha) with 0.1 wt ppm in product for our new regulation. We have 1 reactor (R1) with 1 bed of catalyst (18m3 catalyst in 27m3 reactor). I think we should install one more reactor. But I don't know which case is better between: Case 1: Feed-R1-R2-Stripper-splitter and Case 2: Feed-R1-Stripper-Splitter-R2 (recycle bottom product from splitter to R2)-R1. May you have any advice for our revamp?

Additional info:
Of course that Case 1 is traditional process revamp. But I have just read an article from Chevron, about their process revamp as Case 2. It called SSRS Isocraking (single stage reverse sequencing), licensed by Chevron Lummus Global. In that article, they said that the revamped unit can run at 133% of original design capacity with the existing recycle gas compressor. I think in case 2, R2 is existent reactor and R1 is new one (because R1's volume needs to be bigger than R2) This article named "Hydroprocessing upgrades to meet changing fuels requirement", Jay Parekh and Harjeet Virdi. Unfortunately, It's not for NHT, It's Hydrocracking. Is it O.K if I use Case 2 for my NHT revamp?
(9)
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?
(1)
26/05/2011 Q: We are looking for a simple static bed packed with some adsorbants to remove minor impurities of Mercaptons and Ammonia from an LPG stream. We can use a water trickling Absorber for NH3 removal as suggested by Eric. But we need a similar simple solution like a Sulphur Trap for Mercaptons. One Chinese company claims to have a packed Adsorber. (1)
25/05/2011 Q: One of our clients using LPG as feedstock for Isobutylene extraction is having problems with impurities such as NH3 and RSH though in pppm levels in Feed. How can these be removed? (2)
23/05/2011 Q: Is there any conversion factor between Nm3/hr anf m3/hr?
I am confused because gas flowrate are measured in KNm3/hr while liquified gas flowrate are measured in m3/hr.
(3)
23/05/2011 Q: We are producing about 3m3/hr of lpg to storage. the temperature and pressure in the overhead drum are 31C and 8.5 kg respectively. at same time consuming about 2.94KNm3/hr of lpg in our heater. the pressure of the lpg going to heater is about 1.02kg. What is my net loss or gain of lpg? (1)
22/04/2011 Q: There is a chilling water package that chilled water by use of propane refrigerant cycle. unfortunately propane cycle is polluted by caustic and we decided to wash the lines and equipment. I want to know, is washing with water or low pressure steam is harmful for this? (5)
23/02/2011 Q: What will be the sand filter specification for LPG service? (1)
18/02/2011 Q: Our HP sour gas header battery limit B/V is passing and leads to shut down other supplying unit to replace passing valve. To face this problem in future, maintenance is going to install 2nd block valve after removing originally installed NRV.
Q-1 Will it be successful?
Q-2 Why there are NRVs installed in battery limits for incoming lines?
(2)
08/02/2011 Q: we are facing a problem of hydrate formation in propane BOG (boil off gas) recovery.Th stream is mainly consists of propane ,ethane and low concentration of water (1 ppm volume). How can I determine the solubility of water in propane at low temperature ranging from -40°C to 0°C.
i will be than grateful if can I have the reference papers in the subject.
 
30/01/2011 Q: we face a big problem. We have a caustic unit to sweetened propane and butane in the gas refinery. For caustic regeneration, after oxidation of rich caustic, it converts to De sulfide oil and lean caustic. Wash oil is used to separate De sulfide oil and caustic because of close density of these two product. Sometimes we don't have any wash oil to use and total sulfur at both products increasing awfully. It means that no mercaptane was separated from propane or butane. Do you have any suggestion to separate De sulfide oil from Caustic without wash oil? Do you know any alternative besides wash oil? (2)
18/01/2011 Q: We are facing a high odor rating in Polypropylene. We have merox unit for sulfur removal from LPG & also PRUs. To further enhance sulfur removal (particularly for heavier mercaptans) we have installed Naphtha wash facility in Merox. Still we have odor issue. COS & heavier mercaptans, I think, are the main culprits. Can somebody advise how to improve Merox performance? Or suggest a new facility to cut down sulfur level further? (4)
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.
(2)
26/10/2010 Q: We are fixing for choosing a new compressor for VRU; vapors recovery unit; the light hydrocarbon gas mixture has Mw about 42, K{Cp/Cv} =1.12 , Rc {Pd/Ps} =3.2 , Z factor is about 0.97, Inlet temperature is 43 C and flow rate Q is 600 CFM, discharge Pressure, Pd= 72 psia...
I'd be more than grateful for any tips, Equations, etc to use...
(1)
13/10/2010 Q: Does anybody have any experience of using mixed amines, i.e. DEA+MDEA, for sweetening? I am interested in operational problems like foaming, sludging etc related to mixed amines treatment. (2)
10/10/2010 Q: Our amine system circulation rate is 250 m3/hr. Since commissioning we make up losses by adding fresh amine.
1) Please recommend actions taken to check and monitor health of amine system.
2) What is the typical life of DEA 20 WT % after which whole amine is to be replaced ?
3) Do we have to bleed some amine from reflux drum to reduce system corrosovity?
(1)
07/10/2010 Q: 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. (2)
07/09/2010 Q: Our benzene product tank is internal floating roof tank with N2 blanketing which follow US EPA regulation. However measuring the VOC content at breath out shows as high at 15000 ppm. The internal roof rim seal was replaced and produced only minor improvement.
Is there any plant try to install vapor recovery unit to reduce these emissions? Is there any regulation which requires the benzene tank to be equipment with close system?
(1)
01/09/2010 Q: I would like to know why there are two feed inlets in a Sour Water Stripping tower, but normally only one will be in use and the other will be blinded. In Sour water strippers why is chimney tray provided? (4)
26/05/2010 Q: In our Once Through Hydrocracker, the Fractionator Feed Furnace has options for both Fuel Oil and Fuel Gas Firing. Currently due to some problem in the electrical heater in the Fuel Oil Circuit we are using only fuel gas. Some days back inspection department reported a much higher skin temperature in the radiation section of the Furnace. The same report was also upheld during various cross-checks by other departments. Could this be due to the reason as we are not using Fuel Oil? If so, then could somebody explain? Another thing to consider, we are running at 70% T'Put and design conversion so in general the burners are supposed to operate at the given Heat Duty. (4)
08/03/2010 Q: 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? (3)
19/02/2010 Q: 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. (6)
21/10/2009 Q: Is it safe to consider back pressure of 50-70 kg/cm2g when my PSV set pressure is at 229 kg/cm2g? Why are we limited to 3-5 kg/cm2g back pressure maximum when we are designing the HP flare? API 520 part 1 says that I can consider up to 50% of set pressure of balanced PSV, so can I consider up to 100 kg/cm2 g when my PSV is set at 220 kg/cm2g? If not, then what is the reason? (4)
15/10/2009 Q: I have a PSV with a set pressure of 229kg/cm2g. What could be the back pressure I consider while designing the flare header so that it would be cost effective as well as safe? (4)
06/10/2009 Q: I would like to know the commercial success rate of the following technologies for recovery of LPG and Natural Gas Liquids (NGLs) from Natural Gas:
1. Absortive Process - AET or similar
2. Supersonic Gas Conditioning Process- TWISTER
(1)
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. (2)
24/07/2009 Q: The Delayed Coker Unit (DCU) and the FCC GasCon Dry Gas is treated in an Amine Unit (with MDEA), in order to eliminate H2S, prior to injection into the refinery fuel gas system. However, operational problems have been experienced at the Amine Unit, due to MDEA degradation and the presence of heat stable salts (HSS), among other factors.
We know that HSS formation is due to an irreversible reaction between some contaminants (strong acids anions such as formate, acetate, thiosulfate, thiocyanate and chloride) and the amines molecules. Furthermore, we know that the DCU Gas contains anions such as acetate, formate and cyanide.
However, we have no available information about the contaminant concentration in the DCU Gas or FCC GasCon Dry Gas.
Do you have any information related to a typical contaminant concentration (e.g. strong acids anions) for a DCU and/or FCC GasCon Dry Gas? Moreover, any additional information would be appreciated (E.g. What kind of process do you think would be appropriate for reducing contaminants concentration? We have heard that a water wash stage previous the amine treating could be useful).
(3)
18/07/2009 Q: What is the minimum safe distance between flare stacks and electric high line? Please specify the related code and standard. (2)
04/06/2009 Q: In what situation is a pneumatic test at one kg/cm2 to be preferred to a hydro test at the design pressure of a vessel? (2)
05/05/2009 Q: The context is the following:
- The system is: inlet pipe + control valve + outlet pipe.
- The fluid is natural gas
- The outlet pipeline is buried.
- No outlet pipe insulation.
- The minimum allowable temperature in the outlet pipe is -20°C.
- The minimum temperature at the control valve outlet flange is about -15°C (worst scenario)
The problem is that I need to calculate the length of outlet pipe so that the fluid temperature increase to 0°C.
My data are:
- Outlet pipe material: carbon steel (L360)
- Outlet pipe internal diameter: 570 mm
- Outlet pipe thickness: 20 mm
- Outlet pipe is buried 1 m deep.
- Average air temperature: 11°C
- Wind velocity: 10 m/s
My questions are:
1. Do you know where can find thermal conductivity data for ground? I know it strongly depends on the ground composition but I don't have anything...
2. Could you please share any Excel spreadsheet to perform that calculations?
 
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?
 
10/04/2009 Q: If off gases contain nitrogen and they fired in fired heaters how will it affect NOx levels? (2)
13/03/2009 Q: In a particular complex onshore gas plant, flare network purge is via continuous flow of N2 controlled through flow orifices, purge points being located at the ends of all major headers. There are also a few fuel gas purge connections but these are located close to the flare stack. Under normal operation fuel gas purge points are closed, ie no flow.
I would like to know what would be the risk of stopping all N2 purge gas and starting fuel gas purge. This would lead to the flare network being purged only close to the flare stack. Rest of the network will have to depend on control valves / other vents for a positive gas flow towards the stack.
We can assume for the sake of this discussion that the fuel gas rate is sufficient to safeguard the seal function of preventing air ingress through stack.
(3)
07/03/2009 Q: Is there an international specification for LPG especially with respect to ratio of Propane to Butane. Is there any limitation in using 100% Propane as LPG? (1)
24/01/2009 Q: Are there any methods/systems that can utilize/recover for re-usage gases released into refinery flare system during non emergency periods? (2)
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.
 
07/08/2008 Q: In addressing refinery CO2 management, can you comment on CO2 curtailment from on-purpose hydrogen plants through "minimised" involuntary'steam, internal heat recycle and captive integration?
 
01/08/2008 Q: How can I design sizing of a jet mixer? what are the factors that determine its efficiency? Can a jet mixer also operate with Nitrogen? And how to calculate the consumption of Nitrogen? Is it better than conventional mechanical agitators for highly viscous fluids with congealing nature? (1)
16/07/2008 Q: Is there any commercial process where the reactant is selectively scrubbed to enhance the forward reaction?
Is it possible to dehydrogenate propane / butane to respective olefins without having catalyst deactivation problem?
 
20/05/2008 Q: Does anyone have experience of, or know how to set up a repair testing point for transportation of LPG by rail within the CIS?  
26/04/2008 Q: Just a query about chloride levels reported for produced water
(formation water) from different sources.
Source-1: 1700 mg/L
Source-2: 80,000 mg/L
Source-3: 1,17,000 mg/L
1. There is order of magnitude in difference in chloride levels in
the various sources above. Can someone from comment/advise on wide difference in the values reported? Is it something to do with method of analysis / units?
2. Any guess value for chloride levels in separator gas. As you are aware process simulations does not help to provide such information?
3. In some cases, it is reported in terms formation water and the
values reported are much higher than produced water. Question is whether produced or formation water are not same?
 
27/02/2008 Q: It is a primary requirement of instrument change-over philosophy that all existing field control systems, safety systems and associated field instruments should remain fully operational and functional in their current configuration until new systems are fully installed, tested and and commissioned successfully. The existing field instruments and associated plant control and safety systems will be operating in parallel with the newly expanded facilities until commissioning is successfully completed. My question is that how parallel operation is possible and how the old system is decommissioned.
 
25/02/2008 Q: We have an occurrence where a small topsides ESD is reported to be "passing". Leakage rate has not been quantified. Our performance standard requires topsides ESDVs to provide effective isolation of hydrocarbon inventory on demand. Our assurance routine is to ensure valves close within specification time through system wide ESD test or crediting of unplanned ESD. Valve passing is only picked up through operational experience (eg preparing for intrusive work). Typically, we treat performance standard deviations as safety critical corrective work and ensure mitigating measures are in place until work is completed. Question that has been raised is what leakage rate on topsides SDVs should be considered an unacceptable leakage rate? How do other operators treat incidents like these? What standards are there to check the leakage rate?

(1)
13/02/2008 Q: I am looking for a heat exchanger specialist or a manufacturing company who would be able to help with tube bundle failures which are very regularly occurring on a horizontal thermosyphon reboiler on a sour water stripper.
We are suspecting a mechanical problem like vibration or something else. The tubes are failing in six month to a year even if they are upgraded to stainless steel.
The problem does not seem to be related to corrosion from the process fluid.
(2)
24/01/2008 Q: Is there a process to make Carbon Black Feedstock (CBFS) from natural Gas? (2)
22/01/2008 Q: How much does it cost a refinery and/or petrochemical plant to produce 1 (one) tonne of CO2? I have worked out how much CO2 is produced per barrel of oil, for example, but now want to put a monetary value (or indeed an energy value) on to that tonnage of CO2. Thanks.  
09/01/2008 Q: What could be the cause of gas dryer losing efficiency very fast? Our dryer is molecular sieve dryer.  
09/12/2007 Q: We have a glycol dehydration unit to dry the wet gas by TEG. During the regeneration of TEG, we use natural gas as stripping gas to reach the purity of 99.9wt% of TEG. Note the process under temperature and pressure of 204C and atm.
We have a proposal to substitute the stripping gas with N2 gas.
Please could any one tell me if the N2 will adequate for that process or not ?
 
27/11/2007 Q: How many, and what capacity, Gas To Liquid (GTL) plants are currently operational or under construction? (2)
04/11/2007 Q: What methods are available for the removal or reduction of Phenol from a vacuum heater? (2)
09/10/2007 Q: What are the various processes for Recovery of Sulfur from Acid Gas? (4)
06/09/2007 Q: Please advise on the design and operating temp and pressure for the cryogenic tank of LPG. The composition of LPG is Propane:Butane is 0:100 and 50:50.  
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)